High throughput development method for catalytic hydroprocessing of dirty feedstocks

ABSTRACT

A method for determining a set of operating parameters for a commercial scale plug flow catalytic process and reactor system for hydroprocessing dirty feedstocks, comprises the steps of: feeding selected partial pressures of said feedstock and hydrogen to the inlet the first reactor stage of a first composite multi-stage series-connected laboratory scale plug flow reactor including at least three reactor stages, the catalyst beds of each of said reactor stages including catalyst particles capable of catalyzing the removal by hydrogen of heteroatoms from said heterocyclic molecules; sampling the effluents of each of said reactor stages; measuring the concentration of heterocyclic molecules in said dirty feedstock in the concentrations of heterocyclic molecules and intermediate and final products and by products of the catalytic reaction in the effluents of each of said reactor stages.

FIELD OF INVENTION

This invention relates to methods for the low cost, accelerateddevelopment from discovery to commercial readiness of catalysts and plugflow catalytic processes for hydroprocessing of dirty feedstocks. By theterm “dirty feedstocks” is meant feeds such as C¹⁰-C²⁰⁺ distillates fromcrude oil that contain large heteroatom containing molecules, and cycleoils produced in refinery operations.

BACKGROUND OF THE INVENTION

In order to scale-up a plug flow catalytic process for hydroprocessingof dirty feedstocks, it is necessary to define the impact of feedstockcomposition, time on stream, residence time, catalyst particle size,shape and other characteristics, and temperature profile on reactionrate and selectivity. The first step in a traditional scale-up programgenerally involves the selection, and definition of the intrinsicproperties of, the catalyst. This step is typically performedisothermally with a diluted, crushed or powdered catalyst to minimizemass transfer limitations. A process variable study is performed todetermine the impact of feed composition, space velocity, pressure, andresidence time on reaction rate and selectivity. Activity andselectivity maintenance are then determined over a six to twelve monthoperating period. At the end of the operation, a second process variablestudy is performed to determine whether these properties have changedduring time on stream.

Next, a commercial form of the catalyst is tested in an isothermalreactor. The commercial catalyst is of a larger particle size than thecrushed catalyst and may have a special shape to minimize pressure dropduring operation. The larger particle size generally results in a lowerreaction rate and a selectivity loss due to limitations on mass transferof reactants or products in and out of the catalyst pores. Operationsgenerally consist of performing process variable studies at thebeginning and end of an activity and selectivity maintenance run. Thisoperation can be run in a laboratory scale reactor and typically lastsapproximately one year.

In a commercial trickle-bed hydrotreating reactor, catalyst particleswith diameters of 1.3-3.2 mm and average lengths of 5-10 mm are commonlyused. This range of catalyst sizes reflects the tradeoff between thedesire to provide high catalyst utilization (high catalyst effectivenessfactors) on one hand, and to maintain a manageable pressure drop acrossthe reactor on the other. Within this range of catalyst particle size,the intraparticle diffusion resistance (the diffusion of sulphurcompounds inside the catalyst pores) could significantly reduce theoverall HDS reaction rate. Furthermore, the intra-particle diffusioneffect varies with operating conditions (particularly with temperature)and sulphur compounds of different intrinsic reactivities. While thisissue is very important for a full understanding of HDS mechanisms andprocesses, it has not been adequately and quantitatively addressed inthe open HDS literature, although many HDS studies have been conductedseparately with either full-size catalysts or crushed catalystparticles.

The final step in the scale-up process is to test the commercialcatalyst under adiabatic conditions, normally in a demonstration scalereactor containing one or more reactor tubes. The tubes in thedemonstration scale reactor would have internal diameter ofapproximately 1 inch. In some cases, to further explore heat transfereffects, a configuration containing up to about 6-8 tubes arranged atcommercial spacing could be used. In an exothermic reaction, thetemperature profile depends upon the degree to which heat iscontinuously removed, as in a tubular reactor, or the reactor is simplya fixed bed reactor without a specific heat removal capability. Thetemperature profile can have a significant impact on selectivity,reaction rate, and activity maintenance. The test run also provides ameasure of the tendency for the catalyst to produce hot spots ortemperature runaways. Here again, the operating period can exceed oneyear.

This sequential approach typically takes in excess of three years tocomplete and may not provide all of desired data for scale-up. For manycatalysts, the reaction rate and selectivity may be a function ofresidence time as well as time on stream. This can be the result ofchanges in the catalyst state or form, due to exposure for extendedperiods of time, or it may be due to the changing gas and liquidcomposition from the reactor inlet to the outlet. Examples would includeoxidation from water formed during conversion, formation of a supportover layer, poisoning, e.g., by reaction with hydrogen sulfide andammonia, etc. In addition, surface catalytic reactions and buildup offeed and products in the pores can result in reductions in mass transferrate to the catalyst.

The hydroprocessing of dirty feeds involves a sequential series ofupgrading steps. These include first the removal of a heterocyclicmolecules from the feed by a desulfurization, deoxygenation anddenitrogenation step, followed by a hydro-dearomatization step to getrid of polynuclear aromatics. The catalysts used for these steps must besubstantially immune to sulfur, oxygen and nitrogen poisons and aretypically well-known metal sulfide catalysts. Thereafter, the feed maybe subjected to ring opening step in which carbon-carbon bonds in ringstructures are selectively cleaved to open the rings and some long chainmolecules are broken into shorter segments. This step is also typicallyperformed using a metal sulfide catalyst, however the sulfur levels inthe catalyst are normally much lower than in the catalysts used for theprevious steps. Thereafter, the feed may, if necessary, be subjected toa hydrogenation step for saturating the H—C bonds in the feed. Thecatalyst used in this step is typically a Zeolite supported or bulknon-precious Group 8 metal, although precious group metal catalysts canalso be used. Finally the feed is subjected to a hydrocracking andisomerization step to form saturated hydrocarbons and napthenes of thedesired chain length distribution. The catalysts used in the steptypically comprise a Zeolite supported Group 8 precious metal. The stepsin this process can be performed in separate plug flow reactors.Alternatively, some or all of the steps may be performed in a plug flowreactor having a longitudinally segmented bed, wherein each bed segmentcomprises a catalyst optimized for the given catalytic process step.

Despite the relatively large number of patents covering the differentunsupported catalysts and their applications in hydroprocessing, thereare only a few commercial bulk hydrotreating catalysts. In light of thepotentially huge performance benefits demonstrated in the laboratory,this may look somewhat surprising. However, incorporating a catalystwith extremely high activity in existing refinery process equipment isall but straight forward. In many cases the process and the equipmentwas not designed for the heat release and H2 consumption accompanying avery high activity catalyst. The commercial use of such a catalysts canonly be accomplished by close cooperation between the catalystmanufacturer and the refiner to ensure the expected performance benefitsare achieved without undesired side effects or operational difficulties.In addition to that, the price plays an important role even if theunsupported catalyst contains only Group VIII and Group VI metals. Highconcentration of metals and higher density of unsupported, as comparedto supported, catalysts will increase the reactor fill pricesignificantly. Consequently, there are only limited offerings ofcommercial bulk hydroprocessing catalysts, with one being sold insizable amounts in the present hydroprocessing market. This is theNEBULA1 catalysts family NEBULA-1 and NEBULA-20 are the commercialgrades) of the Albemarle Catalyst company. Clearly there areopportunities for other more advanced catalysts to be offered into themarket, to help further improve product quality and to reduce overallcosts of desulfurization to clean fuels.

Recently, High Throughput Experimentation (HTE) techniques have beenproposed as a source of data for new catalysts and processes. These HTEexperiments are normally performed under conditions that minimize heatand mass transfer effects. Small volumes (less than 2 ml) of catalystand high heat transfer rates are utilized. This approach is useful forcomparing the intrinsic properties of an array of candidate catalystsbut does not provide the data required for scale-up. See, for example,U.S. Pat. Nos. 6,149,882 and 6,869,799.

Combinatorial chemistry is in widespread use in the pharmaceuticalindustry, where it is used to synthesize, purify and evaluate new drugsat a more rapid pace than previously possible. In the field ofpharmaceutical chemistry, the reactions are typically performed at arelatively small scale, since only a small amount of each drug isrequired for testing. Typically, only those drugs which are active inrelevant bioassays are scaled up. The type of chemistry used to generatecommercial quantities of the drugs is rarely the same as that used insmall scale synthesis.

Combinatorial chemistry is being used in petroleum chemistry for thediscovery of new catalytic materials for use in various refiningapplications. However, to have any commercial meaning, one must be ableto correlate the results obtained on a small laboratory scale with thosewhich will be obtained on a commercial scale. One combinatorialchemistry approach used to identify useful catalysts for variousreactions involves placing a plurality of catalysts on a metal plate,contacting the plate with a gaseous reactant, and analyzing the productsobtained via GC/MS. This approach is limited because, at least for anumber of exothermic and endothermic catalytic reactions, it isdifficult to correlate the results obtained on this small scale withthose obtained on a commercial scale. This limitation exists, in part,because the heat transfer obtained on such a small scale cannotreasonably be correlated with what would be observed in a large reactor.There are many recent examples of this combinatorial approach and somerepresentative ones are found in the following references: U.S. Pat. No.5,776,359; U.S. Pat. No. 5,853,356; U.S. Pat. No. 5,939,268; U.S. Pat.No. 6,149,882; U.S. Pat. No. 6,790,322; and U.S. Pat. No. 7,025,828. Inaddition, the issue of process scale up has been addressed to somedegree in U.S. Pat. No. 6,806,087. Here too, the approach is one that isunable to adequately address the full range of scale up issues includingboth heat and mass transfer effects on a commercially relevant scale ofcatalyst operation.

Heat transfer effects are very important in endothermic reactions. Forexample, hydroprocessing such as hydrocracking, hydrodesulfurization,hydroisomerization, which are endothermic reactions, are very sensitiveto heat and mass transfer effects. A small scale reaction which providesan acceptable product mixture may provide an unacceptable level ofsecondary reactions on scale-up due to heat and or related mass transfereffects. Thus, it is quite difficult to extrapolate the results on smallscale endo-thermic reactions to large commercial scale reactors. At thesame time, it is not possible to take advantage of High Throughputcombinatorial chemistry with the use of commercial scale reactors.

It would be advantageous to provide methods for discovering andevaluating optimal catalyst systems using combinatorial chemistrytechniques that are capable of taking the heat and mass transfer effectson product distribution and other reaction characteristics intoconsideration. The present invention provides such methods.

SUMMARY OF THE INVENTION

This invention relates to a low cost, accelerated method for determiningan advantageous combination of reactor structures, catalystcharacteristics, catalyst bed structures and process conditions forscaling up from discovery to commercial readiness a plug-flow catalyticprocess and reactor system for hydroprocessing dirty feed stocks, andhaving high productivity and selectivity to desired products.

The method of the invention involves the use of high throughputlaboratory scale catalytic process development apparatus that includesmultistage series-connected laboratory scale plug-flow reactors toiteratively investigate each of the separate successive hydroprocessingsteps performed on the dirty feed stocks, including removingheteroatoms, saturating polynuclear aromatics, cleaving carbon-carbonbonds in cyclic molecules, saturating unsaturated molecules, andhydrocracking and isomerizing the resulting hydrocarbon molecules. Thecharacteristics and compositions of the effluents of the laboratoryscale reactor stages are sampled and measured during each iteration, andthe results of such measurements made during one iteration are used tohelp determine the choice of catalyst bed characteristics and processconditions in subsequent iterations for improving the productivity andselectivity of hydroprocessing operation.

The process conditions under which the testing is performed includevarious sets of temperatures, pressures, flow rates and relative partialpressures of reactants and reaction products in the catalyst beds of thevarious stages of the laboratory reactors. The testing of the catalystsincludes determining the effects, in the various sets of processconditions, of the relevant catalyst characteristics such as particlesize and shape on the performance of the catalysts in various portionsof the series-connected laboratory scale fixed bed reactors over time.The testing of catalyst bed configurations can include the testing ofthe effects of various configurations of catalyst beds in the laboratoryscale reactor stages, including varying the catalyst characteristicsfrom one laboratory scale reactor stage to the next.

The data generated as a result of this testing enables the design of acommercial scale plug-flow catalytic process and reactor system forhydroprocessing dirty feed stocks in which the catalyst characteristicsand operating parameters, including partial pressures of reactants andproducts, temperatures, pressures and flow rates are optimized. Thus, inaccordance with the method of the invention, the longitudinal gradientsin kinetics, mass transfer and heat transfer characteristics for thevarious reactions occurring within the catalyst beds, or catalyst bedsegments, for each of the upgrading steps for hydroprocessing dirtyfeedstocks in a commercial scale reactor or reactors are investigatedwith the use of composite multistage series-connected laboratory scalefixed bed reactors that effectively permit the segmenting of eachindividual catalyst bed, or bed segment, in the commercial scalereactors into successive longitudinally distributed slices to permit thetaking of measurements to investigate the kinetic, mass transfer andheat transfer characteristics for the different chemistries occurringwithin each of such slices of the catalyst beds or bed segments. Thedata gathered based on these measurements allows the development ofpredictive models using laboratory-scale reactors that describe thebehavior is applicable to large-scale catalytic hydroprocessing systems.

The term “plug flow reactor”, as used herein refers to fixed bedreactors, packed bed reactors, trickle bed reactors and monolithicreactors operating either in a once through or a recycle mode. The term“laboratory scale plug flow reactor” as used herein, refers to a plugflow reactor in which each reactor stage has an internal diameter ofless than 4 inches, preferably less than 2 inches, and more preferablyless than 1 inch; a length of less than 8 feet, preferably less than 4feet, more preferably less than 1 foot; and a catalyst charge of lessthan 800 grams, preferably less than 400 grams, more preferably lessthan 25 grams (excluding inert diluent particles charged to thereactor).

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a schematic representation of a composite multistage,series-connected, plug flow reactor in accordance with the invention;

FIG. 2 is a schematic representation of a composite multistage,series-connected, plug flow reactor and a parallel multistage,series-connected, probe reactor in accordance with the invention;

FIG. 3 is a schematic representation of a composite multistage,series-connected, plug flow reactor and a fluid dynamically linked,single stage probe reactor in accordance with another embodiment of theinvention;

FIG. 4 is a schematic representation of a composite multistage,series-connected, plug flow reactor and a fluid dynamically linked,multistage, series-connected, probe reactor in accordance with theinvention;

FIG. 5 is a schematic representation of a multistage, compositeseries-connected, plug flow reactor disposed in a constant temperatureenvironment in the form of a fluidized sand bath in accordance theinvention;

FIG. 6 is a schematic representation of a plurality of compositemultistage, series-connected, fixed bed reactors disposed in the commonfluidized sand bath in accordance with the invention;

FIG. 7 is a schematic representation of a plurality of compositemultistage, series-connected, plug flow reactors configured to receivecontrolled variable inputs in accordance with the invention;

FIG. 8 is a graph useful for determining the Thiele Modulus of acatalyst;

FIG. 9 is a graph of the Effectiveness Factor versus Thiele Modulus fora catalyst;

FIG. 10 is a graph of Effectiveness Factor versus conversion for crushedand commercial scale catalysts;

FIG. 11 is a schematic representation of a plug flow reactor arrangementin accordance with another embodiment of the invention;

FIG. 12 is a schematic representation of a multistage, compositeseries-connected, isothermal plug flow reactor in accordance with theinvention;

FIG. 13 illustrates an assembled, schematic diagram of reactors and aseparator in accordance with one embodiment of the present invention;

FIG. 14 illustrates an assembled, schematic diagram of the reactors andthe separator in accordance with another embodiment of the presentinvention; and

FIG. 15 illustrates an assembled, schematic diagram of the reactor andthe separator in accordance with yet another embodiment of the presentinvention.

DESCRIPTION OF THE PREFERRED EMBODIMENTS

Referring to FIG. 1 of the drawings, the composite multistage laboratoryscale plug flow reactor 11 of a first embodiment of the invention ismade up of three series-connected plug flow stages 13, 15 and 17, inthis case fixed bed stages, each of which contains a bed of catalystparticles 19, 21 and 23. A sampling valve 25 is connected between theoutput of the first reactor stage 13 and the input to the second reactorstage 15 and has an output 26 for sampling the effluent from the firstreactor stage 13 for analysis. A sampling valve 27 is connected betweenthe output of the second fixed bed reactor stage 15 and the input to thethird fixed bed reactor stage 17 and has an output 28 for sampling theeffluent from the second reactor stage 15 for analysis. A sampling valve29 is connected to the output of the third fixed bed reactor stage 17and has an output 30 for sampling the effluent of the third reactorstage 17 for analysis. The output of the third reactor stage 17 isconnected through the valve 29 to, e.g., a product accumulator (notshown). The feed to the multistage fixed bed reactor 11, which normallyis fresh reactant feed, is connected to the inlet of the first fixed bedreactor stage 13 from a source 31. A sampling valve may also beinstalled in the line between the feed source 31 and the inlet to thefirst fixed bed reactor stage 13 in order to permit analysis of thefeed.

The multistage fixed bed reactor 11 is contained in a temperaturecontrol device 33 that, for an exothermic reaction such as theFischer-Tropsch reaction, could contain a material, such as circulatingboiling water or a fluidized sand bath, for extracting heat from thereactor 11 in order to maintain the multistage reactor 11 at asubstantially constant temperature. For an endothermic process, such asparaffin dehydrogenation or catalytic reforming, the temperature controldevice 33 could contain apparatus, such as an electrical heater, tosupply heat to the fixed bed reactor 11 in order to maintain thesubstantially constant desired temperature. Alternatively, for bothexothermic and endothermic catalytic processes, the temperature controldevice 33 can consist of a fluidized sand bath heater in which themultistage reactors are immersed.

Each of the catalyst beds 19, 21, and 23 in the reactor stages of themultistage reactor 11 replicates a longitudinal portion of the catalystbed of a large fixed bed reactor and permits the measurement andanalysis of the characteristics and performance of successivelongitudinal portions of a large catalyst bed, thereby allowingdetermination of longitudinal gradients in reactor characteristics andperformance that heretofore have been inaccessible. While reactor 11 hasbeen shown as having three series-connected stages, it is equallypossible to have a larger number of series-connected stages, e.g., fouror six stages, in order to analyze the performance of the compositecatalyst bed at a greater number of points along its length.

Depending on the reaction being studied and the data needed, theanalysis of the feed and the effluent from the reactor stages caninclude, e.g., conventional GC/MS or UV or IR characterization of thereactant and product stream(s), and/or analysis of the catalyst systemby XRD, diffuse reflectance IR or other spectroscopic techniques thatare well known in the art. These studies would allow the performanceattributes of the system to be quantified as a function of thelongitudinal position in the catalyst bed. Such knowledge would allowthe system to be optimized with direct knowledge of the catalyticreaction kinetics and performance attributes of each point and permitthe design of catalyst systems in which, e.g., the catalyst particlesmay have different chemical or physical characteristics in differentportions of the catalyst bed so as to operate at peak productivity orselectivity as a function of the local environment.

The catalyst beds in the reactor stages 13, 15 and 17 may be a crushedor powdered catalyst or a commercial-size catalyst. Most measurementsmade in gathering data for the scale up of a catalytic reactor need tobe made with the reactor operating in a substantially isothermal regime.In order for the reactor stages 13, 15 and 17 to operate in asubstantially isothermal regime, the catalysts in the beds 19, 21 and 23are diluted with an inert particulate matter, typically in a ratio of upto about 8-10 to 1. For measurements being made with the reactoroperating in a substantially adiabatic regime, the catalyst in the beds19, 21 and 23 is less diluted, and depends on the heat of reaction ofthe process under study and reactor diameter. The ratio of catalystparticles to diluent particles in a catalyst bed depends upon a numberof factors, including the amount of heat generated by the reaction andthe activity of the catalyst particles in the bed. The appropriate ratiofor a given reaction, catalyst, reactor diameter and catalyst particlesize can easily be determined by one of ordinary skill in the art by asimple experiment.

A commercial-size catalyst in a fixed bed reactor typically has particlesize of about 1 to 5 mm. the catalyst particles can be in any one ormore a variety of shapes, e.g., round, tubular, trilobe, toroidal, etc.The crushed or powdered catalyst, which is typically formed by crushinga commercial-size catalyst, typically has a particle size of about0.10-0.20 mm. the crushed or powdered catalyst particles are normallypreferably as small as can be obtained while still retaining aperformance qualities of the catalyst. The interior diameter of areactor stage should be about 10 times the diameter of the smaller ofthe diluent or catalyst particles and the minimum would typically be inthe range of about 10 to 50 mm (0.4 to 2 inches) for a bed containingcommercial-size catalyst particles and diluent. Crushed or powderedcatalyst particles are typically more active than the commercial-sizecatalyst particles because of lower mass transfer resistance. Therefore,in order for a reactor containing a bed of crushed or powdered catalystto operate at the same temperature as a similar reactor containingcommercial-size catalyst, the ratio of inert diluent particles tocatalyst particles in the bed of crushed or powdered catalyst particlesnormally needs to be higher than that of the bed containingcommercial-size catalyst particles in order that the heat release perunit volume of the to catalyst beds is the same.

The interior diameter of a reactor containing crushed catalyst, can, ifdesired, be smaller, in the range of about 5 to 12 mm, than that of areactor containing the commercial size catalyst. For reasons offlexibility in the use of the multistage reactor 11 in differentapplications, however, it may be preferable that the crushed catalystbed have the same interior diameter as that required for a bedcontaining commercial-size catalyst particles. Alternatively, theinterior diameter of a reactor being used with a bed of crushed orpowdered catalyst particles may be reduced by the use of a thermallyconductive sleeve within the reactor.

The minimum height of a reactor stage is determined either by mixing orheat release considerations. For isothermal operation, if mixing is thelimiting factor, the height should be sufficient to avoid bypassing.Typically, this would be at least about 50 times the average diameter ofthe particles, or about 50 to 250 mm (2 to 10 inches) for a reactorstage containing a bed of commercial-size catalyst particles. Becausethe feed is progressively converted as it traverses the stage of themultistage reactor 11, the concentration of fresh feed in the successivereactor stages decreases from one stage to the next. If it is desired tohave constant conversion in each reactor stage, the lengths of thecatalyst beds 19-23 can be progressively longer in each of thesuccessive reactor stages 13-17. If the reactor 11 is to operate in theadiabatic regime, one would tend to use a lower ratio of inert diluentand a larger diameter reactor.

Referring to FIG. 2 of the drawings, there is illustrated a secondembodiment of the invention in which elements that are the same as inthe embodiment illustrated in FIG. 1 are numbered similarly. This secondembodiment includes a composite multistage reactor 11 that is the sameas the multistage reactor 11 of FIG. 1. A composite multistage probereactor 35, in which each reactor stage can be the same as thecorresponding reactor stage of multistage reactor 11, is operated inparallel with the multistage reactor 11. Both of the multistage reactor11 and the probe reactor 35 are contained in a temperature controldevice 33 that can be the same as the types discussed above. If desired,the probe reactor 35 can be contained in a temperature control deviceseparate from the temperature control device 33 in which the reactor 11is contained, thereby permitting the operation of the probe reactor 35at a temperature different from that of the multistage reactor 11.

The composite reactor 35 has three series-connected reactor stages 37,39, and 41 that contain catalyst beds 43, 45 and 47, respectively. Asampling valve 49 is connected between the output of probe reactor stage37 in the inlet of the probe reactor stage 39 and has an output 50 forsampling the effluent from reactor stage 37. A sampling valve 51 isconnected between the output of reactor stage 39 and the input ofreactor stage 41 and has an output 52 for sampling the effluent from thereactor stage 39. A sampling valve 53 is connected between the output ofreactor stage 41 and, e.g., a product accumulator (not shown), and hasan output 54 for sampling the effluent from reactor stage 41. The freshreactant feed from source 31 is connected to the inlet of the firstprobe reactor stage 37. A control and sampling valve can be connectedbetween the source 31 in the inlet to the first probe reactor stage 37for selectively controlling the amount of feed to the probe reactor andto permit the sampling of the feed for analysis. Also connected to theinlet to the first probe reactor stage 37 is a source 55 of a materialto be controllably added to the input of the first probe reactor stage37 for ascertaining the effect of such addition on the characteristicsand performance of the stages of the probe reactor 35. A source 57 isconnected to the inlet of the second probe reactor stage 39 forselectively adding a material to the input of such a second probereactor stage for ascertaining the effect of such addition on thecharacteristics and performance of the second and third probe reactorstages 39 and 41. A source 59 is connected to the input of the thirdprobe reactor stage 41 for selectively adding a material to the input ofsuch probe reactor stage for ascertaining the effect of such addition onthe characteristics and performance of the third probe reactor stage 41.In this embodiment of the invention, the catalyst beds 43, 45 and 47 ofthe probe reactor 35 are preferably the same as the catalyst beds 19, 21and 23 of the multistage reactor 11, respectively.

The use of the composite multistage probe reactor 35 allows one tomeasure the transient response of the system to permanent or temporarychanges in the feed composition at any stage of the multistage reactor11 by comparing the characteristics and performance of the relevantstages of the probe reactor 35 over time in response to the change ininput with the characteristics and performance of the correspondingstages of the multistage reactor 11. Introduction of a change in gas orliquid input to the third reactor stage of the probe reactor 35 allowsone to measure the impact of the changed component on the reaction rateand selectivity of the third reactor stage catalyst bed of themultistage reactor 11 with time. Introduction of the change to thesecond probe reactor stage allows one to measure the impact on thesecond and third stage catalyst beds of the multistage reactor 11. Thisis equivalent to measuring the response to a change in conditions of anysmall segment of the catalyst bed in a commercial-size fixed bedreactor. For example, raising the gas feed rate to any reactor stage ofthe probe reactor 35 by having one of the sources 55, 57 or 59 andadditional fresh feed into the stage of the probe reactor 35 to which itis connected, would allow the investigation of the changes inincremental performance of that stage and following stages resultingfrom the change in input over time.

It is also possible to use the sources 55, 57 or 59 to vary theconcentrations of the trace components present in the fresh feed in aselected probe reactor stage, for instance by adding fresh reactant feedhaving a higher or lower concentration of such trace components, inorder to quantify the effect of such trace components on various partsof the composite catalyst bed under a full range of operatingconditions. By doing this it would be possible to map the criticallongitudinal portions of the composite catalyst bed in a commercialsystem in which the catalyst is most vulnerable to poisoning or otherinhibitory reactions caused by poisons or other natural byproducts ofthe reaction being practiced. The probe reactor 35, and other versionsof probe reactor as discussed below with relation to other Figures, canalso be used to investigate the transient response of a reactor totemporary changes in the composition of the feed or prior stage effluentto various points in a composite catalyst bed by temporarily adding thematerials of interest to a selected stage of the probe reactor 35 andmonitoring the time dependent response of that stage and followingstages of the probe reactor 35 to such added materials both during andafter the time that such materials are added.

Referring to FIG. 3 of the drawings, there is illustrated anotherembodiment of the invention in which elements that are the same as inthe embodiments of FIG. 1 are numbered similarly. In this embodiment ofthe invention, the probe reactor 101 can consist of a single plug flowreactor stage whose inlet is selectively fluid dynamically linked to aselected stage of the composite multistage fixed bed reactor 11. Otherconfigurations for the single stage probe reactor 101 are discussedbelow. The valve 103 is connected between the output of the firstreactor stage 13 and the input of the second reactor stage 15 of themultistage reactor 11 and has outputs 105 and 107 for selectivelysampling of the effluent of the reactor stage 13 and selectivelyconnecting a portion of the effluent of the reactor stage 13 to theinput of the probe reactor 101, respectively. The valve 109 is connectedbetween the output of the reactor stage 15 and the input to the reactorstage 17 of the multistage reactor 11 and has outputs 111 and 113 forselectively sampling of the effluent of reactor stage 15 and selectivelyconnecting a portion of the effluent of reactor stage 15 to the input ofprobe reactor 101, respectively. The valve 107 is connected between theoutput of reactor stage 15 and a product accumulator and has outputs 117and 119 for selectively sampling of the effluent of reactor stage 15 andselectively connecting a portion of the effluent of reactor stage 15 tothe input of probe reactor 101, respectively. The probe reactor 101 alsoreceives inputs from the feed source 31 and from a source 121. The probereactor 101 and the catalyst bed contained therein in this embodiment ofthe invention is preferably the same as the reactor stage and catalystbed contained therein in the multistage reactor 11 following the onehaving a portion of its effluent connected to the input of the probereactor 101. The single stage probe reactor may, for example, be used toperform the same investigations as were described above with relation tothe multistage probe reactor embodiment of FIG. 2.

Referring to FIG. 4 of the drawings, there is illustrated anotherembodiment of the invention in which elements that are common to theembodiments of FIGS. 1 and 2 are numbered similarly. In this embodimentof the invention, the probe reactor 35 consists of a compositemultistage series-connected plug flow reactor in which the reactorstages may be the same as the multistage series-connected probe reactor35 depicted in FIG. 2 of the drawings. In this embodiment, however, thestages of the probe reactor 35 are selectively fluid dynamically linkedto selected stages of the composites multistage series-connected reactor11 by selectively connecting a portion of the effluent of one or morestages of the composite multistage series-connected reactor 11 to one ormore selected stages of the probe reactor 35. The valve 123 is connectedbetween the output of the first reactor stage 13 and the input of thesecond reactor stage 15 of the multistage reactor 11 and has outputs 125and 127 for selectively sampling the effluent of the first reactor stage13 and connecting a selected portion of the effluent of reactor stage 13to the inlet of probe reactor stage 39, respectively. The valve 129 isconnected between the output of reactor stage 15 and the input toreactor stage 17 of the multistage reactor 11 and has outputs 131 and133 for selectively sampling the effluent of reactor stage 15 andselectively connecting a portion of the effluent of reactor stage 15 tothe input of probe reactor stage 41, respectively. The fresh reactantfeed from the source 31 is connected to the input of the first probereactor stage 37. Control and sampling valves (not shown) may beconnected in the line between the fresh reactant feed and the probereactor stage 37 to control the amount of fresh reactant feed suppliedto the probe reactor 35 and to permit the analysis of its content. Alsoconnected to the input to the first probe reactor stage 37 is a source55 of a material to the selectively added to the input of the firstprobe reactor stage 37 for ascertaining the effect of such addition tothe stages of the probe reactor 35. A source 57 is connected to theinput of the second program per stage 39 for selectively adding amaterial to the input of such a second program per stage forascertaining the effect of such addition on the second and third probereactor stages 39 and 41. A source 59 is connected to the input of thethird probe reactor stage 41 for selectively adding a material to theinput of such probe reactor stage for ascertaining the effect of suchaddition on the third probe reactor stage 41. In this embodiment of theinvention, the catalyst beds 43, 45 and 47 of the probe reactor 35 arepreferably the same as the catalyst beds 19, 21 and 23 of the multistagereactor 11, respectively.

Referring again to FIG. 3 of the drawings, the probe reactor 101 canconsist of a substantially fully back-mixed reactor instead of a singlestage fixed bed reactor stage 101, such as discussed above. Thedistribution a catalyst, feed and products in the back-mixed probereactor 101 a substantially uniform and so, if the probe reactor 101receives only effluent from a stage of reactor 11, it corresponds to asingle, narrow, horizontal slice at the inlet of the catalyst bed of thestage of multistage reactor 11 following the stage that has a portion ofits effluent connected to the input of the probe reactor 101. Bycontrolling the relative concentrations of fixed bed reactor stageeffluent and fresh feed, it will is possible for the back-mixed probereactor to simulate any selected horizontal slice of the fixed bedreactor stage whose effluent is connected to the back-mixed probereactor. The back-mixed probe reactor 101 can, for instance, be atwo-phase fluidized bed reactor, a three-phase slurry reactor, or athree phase ebulated bed reactor.

Alternatively, the probe reactor 101, instead of being a fullyback-mixed reactor such as discussed above, can be a two-dimensionalcatalyst array such as disclosed in Y. Jiang et al, Chemical EngineeringScience, vol 54, pp 2409-2419 (1999). Such a probe reactor can be usedto investigate the intrinsic characteristics of a plurality of crushedcatalysts in the presence of different mixes of feed, effluent andproduct.

In the embodiments of FIGS. 2, 3 and 4 of the drawings, stages of theprobe reactor 101 and 35 receive as inputs combinations of controlledamounts of one or more of the fresh reactant feed, effluent from aselected stage of the multistage reactor 11 and other feeds. Such otherfeeds may, for instance, consist of additional fresh reactant feed,additional product gases or liquids produced during the reaction takingplace in the composite multistage reactor 11, or contaminants that maybe present in the fresh feed used during operation of a commercialreactor.

The reactant and other material feeds, and reaction products andbyproducts in reactor effluents supplied or generated in the embodimentsof the invention described herein may be either gaseous, liquid or mixedphase (such as e.g., gas/liquid or two or more immiscible liquids).Feeds and effluents consisting of gases can be handled using well knownconventional back pressure regulators and gas flow control systems withmass flow controllers. Controlled amounts of liquids can be pumped inhigh-pressure environments using known pumps such as a Ruska pump or aSyringe pump. If the effluent from a reactor stage or the feed containsmultiple phases, particularly if such phases are immiscible, such aswater and hydrocarbons or liquid and gas, it is important to avoid slugflow. In such case, sampling valves may consist of e.g., iso-kineticsampling valves such as available from Prosery AS, or splitters such asdescribed in U.S. Pat. No. 4,035,168. Alternatively, the stream may besampled immediately after a static mixer such as available from ProseryAS, which homogenizes the multiphase stream. In combining immisciblefeeds or feeds and effluent to a reactor stage, or in conducting themultiphase effluent from the outlet of one reactor stage to the inlet ofthe following reactor stage in a series-connected multistage reactor, itis typically the practice to manifold of the streams into a line havinga high Reynolds number similar in concept to a fuel injection system inan automobile engine. Alternatively, static mixers such as availablefrom Prosery AS or from Admix, Inc., Manchester, N.H., can also be used.In this case, some simple initial testing may be desirable to confirmthat the operating conditions are leading to the homogeneity of thestream passing through the device. If the gas and liquid are well mixedin a transfer line, it is possible, for instance, to take a combinedliquid and gas sample in a sample bomb connected to the reactor line viadouble block valves. The bomb would be at atmospheric pressure orslightly above. The block valves would be opened and liquid and gaswould be allowed to flow into the bomb. The two block valves would thenbe closed, the sample bomb removed from the reactor and the contentsanalyzed. The presence of a small concentration of an inert gas such asArgon in the stream can be used to allow closure of the materialbalance. Alternatively, if the phases are not well mixed, one couldemploy gas/liquid separators and analyze the gas and liquid phasesseparately with an internal standard such as He or Ar and overall carbonbalance analysis to link the two. This could be accomplished e.g., byusing a gas sample bomb attached to the top of the line and a liquidsample bomb attached to the bottom of the line. A major area of concernin understanding and controlling the characteristics and performance ofa plug flow reactor is the adsorption or reaction of a feed component,product or byproduct with the catalyst surface. For instance, in ahydrocracking processes, materials such as ammonia, carbon monoxide,hydrogen sulfide, can tie up active catalyst sites, reduce reaction rateand adversely impact product selectivity. The reactions caused by thesematerials can take time to equilibrate and can also take time to bereleased after removal of the material from the feed stream to thereactor.

Ammonia is known to react with hydrocracking catalysts, causing activityto decline and line out. Upon removal of the ammonia from the feed,hydrogen can be used to remove the ammonia from the catalyst surface. Ininvestigating the effects of ammonia on different portions of thecomposite catalyst bed, ammonia can be added to the inlet of any of thestages of probe reactor, thereby replicating the effect of the presenceof ammonia in the feed to a selected longitudinal slice of the compositecatalyst bed. By controlling the conversion level in a given catalystslice, e.g., by adjusting temperature and/or flow rate and/or reactantpartial pressures in a probe reactor stage, it is possible to define theeffect of the ammonia under various operating conditions. By varying thehydrogen concentration in the feed to one or more probe reactor stages,it is possible, for example, to investigate the effect of increasedhydrogen on the ammonia-contaminated catalyst in different portions ofthe composite catalyst bed, e.g., the bed with the greatest activitydecline.

Carbon monoxide is tightly held on many Group VIII metal catalysts,which can reduce available surface for hydrogen, thereby making hydrogenactivation the rate limiting step. By varying the concentrations ofcarbon monoxide (or carbon dioxide) and hydrogen in the feed to selectedstages of the probe reactor 35 or 101 and comparing performance of therelevant probe reactor stages with the corresponding stages of themultistage reactor 11, it is possible to determine the impact of carbonmonoxide and hydrogen concentration on reaction rate and selectivity.The use of a multi-stage probe reactor allows for testing of the impactat various conversion levels by e.g., by adjusting temperature and/orflow rate and/or reactant partial pressures.

The addition of water to a plug flow reactor in heavy oil upgrading andconversion processes has been shown to have a positive impact under someconditions and a negative impact under others. Adding controlled amountsof water (or other oxygenated molecules) to selected stages of the probereactor 35 or 101 would permit the study of the impact of the addedwater on reaction rate and selectivity in selected longitudinal slicesof the composite catalyst bed by comparing the characteristics andperformance of the relevant stages of the probe reactor with thecorresponding stages of the multistage reactor 11.

The amount of Conradson carbon is usually utilized in correlations forhydrotreater performance. In general, carbon and heavy wax deposits on acatalyst inhibit the diffusion of reactants to the catalyst surface andthe removal of reaction products from the catalyst surface. This tendsto lead to activity reduction via unwanted side reactions with depositson the catalyst surface or with the diffusion limited reactants or both.In the case of beds containing commercial-size catalyst particles wherethe diffusion path is the longest, this sort of diffusion limitation canlimit overall catalyst life and require costly steps to maintain systemperformance Adding different molecular weight fractions of thesematerials to a selected stage of the probe reactor 35 or 101 would allowthe determination of what portion of the composite catalyst bed isimpacted the most. The effects of various regeneration techniques suchas by the addition of hydrogen, water, or a light solvent can also bedetermined my controlling the feeds to the relevant stages of the probereactor 35 or 101, thereby to define the preferred rejuvenationtechnique. These issues will be particularly important in processing ofheavy sulfur containing feeds from tar sands, shale, heavy oil deposits,and coal. These feeds are known to carry many contaminants that can leadto catalyst poisoning, and in situ regeneration, in order to avoid thecost of frequent replacement with fresh unused catalyst, is frequentlythe only means to make the overall process economically viable.

Polynuclear aromatics are also known to inhibit a catalyst by formingcarbonaceous overlayers on catalyst sites that reduce selectivity andactivity of hydroprocessing catalysts. The effect of the presence ofpolynuclear aromatics in the feed at various longitudinal portions of acomposite catalyst bed of a plug flow reactor can be determined byadding the polynuclear aromatics to a selected stage of the probereactor 35 or 101 and comparing the characteristics and performance ofthe relevant stages of the probe reactor with the corresponding stagesof the multistage reactor 11. This can be used to help define in whatportion of the composite catalyst bed the polynuclear aromatics havetheir greatest impact, and what can be done to improve the processdesign and catalyst performance, by comparison of the performance andcharacteristics of the relevant stages of the probe reactor 35 or 101with the corresponding stages of the multistage reactor 11.

Referring again to FIG. 3 of the drawings, the probe reactor 101 canconsist of a substantially fully back-mixed reactor instead of a singlestage plug flow reactor stage 101, such as discussed above. Thedistribution a catalyst, feed and products in the back-mixed probereactor 101 a substantially uniform and so, if the probe reactor 101receives only effluent from a stage of reactor 11, it corresponds to asingle, narrow, horizontal slice at the inlet of the catalyst bed of thestage of multistage reactor 11 following the stage that has a portion ofits effluent connected to the input of the probe reactor 101. Bycontrolling the relative concentrations of fixed bed reactor stageeffluent and fresh feed, it will is possible for the back-mixed probereactor to simulate any selected horizontal slice of the fixed bedreactor stage whose effluent is connected to the back-mixed probereactor. The back-mixed probe reactor 101 can, for instance, be atwo-phase fluidized bed reactor, a three-phase slurry reactor, or athree phase ebulated bed reactor.

Referring now to FIG. 5 of the drawings, the series-connected reactorstages of the composite multistage laboratory scale plug flow reactor,in this case a fixed bed reactor, according to the invention can bearranged in parallel with one another in a temperature control devicefor a more compact and convenient configuration. In this arrangement thecomposite multistage reactor 501 is made up of three series connectedreactor stages 503, 505 and 507 which are disposed in temperaturecontrol device constituted by a heated or cooled fluidized sand bath509. The reactant feed gas is preferably connected from a source file athrough a preheat coil 513, which is also disposed in the fluidized sandbath 509, to the inlet of the first reactor stage 503. Liquid, which maybe a reactant, is fed from the feed pump 515 through the preheat coil513 to the inlet of reactor 503. Sampling valves may be connected in theboth the gas and liquid feed lines for sampling the gas and liquidfeeds. The preheat coil 513 is used to heat the gas and liquid feeds tothe appropriate temperature for being supplied to the multistage reactor501. The outlet of reactor 501 is connected to the inlet of reactor 505through a sampling valve 517. The outlet of reactor stage 505 isconnected to the inlet of reactor stage 507 through the sampling valve519, and the outlet of reactor stage 507 is connected through a samplingvalve 521 to the separator 523. Each of the sampling valves 517, 519 and521 have an outlet selectively connected to a probe reactor 523 forsupplying effluent to the probe actor 523. Each of the sampling valves517, 519 and 521 also has an outlet to permit sampling of the effluentfrom the respective reactor stage to whose output of the sampling valveis connected.

Referring now to FIG. 6 of the drawings, there is illustratedschematically, an arrangement of three composite multistageseries-connected laboratory scale plug flow reactors 531, 533 and 535and arranged in a fluidized sand bath 537. The stages of each of themultistage reactors are arranged in parallel with one another in thesame manner as the stages of the reactor 501 in FIG. 5, and each of themultistage reactors 531, 533 and 534 is preferably preceded by a preheatcoil that can be the same as the preheat coil 513 illustrated in FIG. 5of the drawings. A single-stage probe reactor 538, which can be any ofthe types described above with relation to probe reactor 101 of FIG. 3of the drawings, is arranged between the series-connected reactors 533and 535 in the fluidized sand bath 537 and selectively receives inputsof either the reactant feed or the effluent of any of the reactor stagesof the series connected reactors 533 and 535 by means of sampling valves(not shown) that may be the same as the sampling valves 517, 519 and 521illustrated in FIG. 5 of the drawings. Each of the reactors 531, 533,538 and 535 receives reactant feed from sources 539, 541, 543, and 545,respectively, that can be all the same feed source. The outlets of thelast stages of each of the reactors 531, 533, 538, and 535 arepreferably connected to the separators or product accumulators 547, 549,551 and 553, respectively, which all may be constituted by a singleseparator or product accumulator.

The arrangements of FIGS. 5 and 6 have the advantage that the fluidizedsand bath need not be so deep as it would be if the reactors werearranged vertically, and in that the sampling valves 517, 519 and 521can be situated above the fluidized sand bath and so are accessible formaintenance or adjustment during operation of the multistage reactors.If the effluents from the stages of the multistage reactors containmultiple phases, the transfer lines connecting the outlet of one reactorstage to the inlet of the following reactor stage need to be configuredin such a way as to avoid a slug flow in the lines. As described above,this can be accomplished using lines having high Reynolds numbers orwith the use of static mixers. The sampling valves 517, 519 and 521 canbe iso-kinetic sampling valves, although other arrangements such asdescribed above can also be used. Additionally, the conduits connectingthe outlet of one reactor stage to the inlet of the followingseries-connected reactor stage are designed for non-slugging flow, forinstance by using static mixers.

Having a plurality of composite multistage series-connected reactorsdisposed in a common temperature environment, such as constituted by thefluidized sand bath 537, or as described above with relation to FIGS. 2through 4 of the drawings, permits the simultaneous investigation ofvarious characteristics of a catalytic process for substantiallyaccelerating the scaling up of the reaction to commercial application.For instance, using the configuration of FIGS. 5 and 6 as an example, ifthe multistage reactor 535 contains crushed catalyst particles dilutedwith an inert diluent for isothermal operation, and the reactor 533contains commercial scale catalyst particles also diluted with an inertdiluent for isothermal operation, and the reactor 531 containscommercial scale catalyst particles in a concentration suitable foradiabatic operation, the kinetic, mass transfer and heat transfercharacteristics of the catalytic process can be investigatedsimultaneously in the isothermal reactors, and the resulting reactormodel derived from the data obtained from the isothermal reactors can beconfirmed by the data obtained from the adiabatic reactor.

Other experiments to be performed that aid in the scaling up of acatalytic process include, for example, investigating thecharacteristics of a plurality of different catalysts simultaneously.Alternatively, a crushed catalyst in the catalyst beds of one multistageseries-connected reactor could be compared with a plurality of differentshapes or sizes of commercial-size versions of the catalyst in thecatalyst beds of other multistage series-connected reactors, alldisposed in a common constant temperature department. In an alternativearrangement, it is also possible to have different catalysts indifferent reactor stages of the multistage series-connected reactor 11for testing the catalysts in series. Using such an arrangement, one candesign a layered composite catalyst bed in which the intrinsic behaviorof each catalyst layer is matched to the local kinetic and mass transferenvironment, so that the overall response of the system is variedlongitudinally so as to obtain behavior characteristics in eachlongitudinal portion of the composite reactor that are optimum forprocess performance If a plurality of multistage series-connected plugflow reactors is disposed in separate, independently controllabletemperature control devices, a plurality of different heat removallevels can be investigated in parallel.

Referring now to FIG. 7 of the drawings, the module 151 contains aplurality of parallel laboratory scale plug flow reactor stages 151-1through 151-n. The module 151 includes a temperature control device 152surrounding the module 151 for controlling the temperature of theambient experienced by the reactor stages 151-1 through 151-n. In thecase of an exothermic reaction, such as the Fischer-Tropsch reaction,the temperature control device may consist of an enclosure containingcirculating boiling water for extracting heat from the reactor stages151-1 through 151-n. For an endothermic process, such as e.g.,dehydrocycloaromatization, steam reforming or hydroprocessing, thetemperature control device can comprise apparatus, such as an electricalheater, for supplying heat to the reactor stages 151-1 through 151-n.For either exothermic or endothermic reactions, the temperature controldevice 152 may alternatively comprise a fluidized sand bath heater inwhich the multistage reactors are immersed.

Each of the reactor stages 151-1 through 151-n contain a catalyst bed153-1 through 153-n. The modules 155 and 157 can be identical to themodule 151, and contain a plurality of parallel plug flow reactor stages155-1 through 155-n and 157-1 through 157-n, respectively. Each of theparallel reactor stages in the modules 155 and 157 contain catalyst beds159-1 through 159-n and 161-1 through 161-n, respectively. In theillustrated embodiment, the outlet of each of the reactor stages inmodule 151 is connected to the inlet of the corresponding reactor stagein module 155, and the outlet of each of the reactor stages in module155 is connected to the inlet of the corresponding reactor stage inmodule 157. Thus, the series connected reactors stages 151-1, 155-1 and157-1 form a composite multistage series-connected fixed bed reactor.Similarly, the other sets of series connected reactor stages in themodules 151, 155 and 157 also form composite multistage series-connectedfixed bed reactors. The modules 151, 155 and 157 may contain any desirednumber of parallel reactor stages depending upon the application. Forinstance, each module might contain four or eight or even 16 parallelreactor stages. Is also possible to have additional modules of parallelreactors stages, with each of said parallel reactors stages beingconnected in series with the corresponding reactor stages of thepreceding and succeeding modules. For instance, there might be four orsix modules in a given application.

The modules 155 and 157 are surrounded by temperature control devices158 and 160, respectively, that may be the same as, or common with, thetemperature control device 152 that surrounds the module 151. Samplingvalves 163-1 through 163-n are connected between the outlet of eachreactor stage in the module 151 and the inlet of the correspondingreactor stage in module 155. Sampling valves at 165-1 through 165-n areconnected between the outlets of each of the reactor stages in module155 in the inlet of the corresponding reactor stage in module 157. Freshreactant feed is fed from a source 167 through control valves 169-1through 169-n to the inlets of each of the reactor stages 151-1 through151-n of module 151 for supplying controlled amounts of reactant feed tothe inlets of the respective reactor stages. The plug flow reactor 171also receives fresh reactant feed gas from the source 167 at its inlet,and has its outlet connected to the inlets of the reactor stages 151-1through 151-n through control valves 173-1 through 173-n, respectively,for supplying controlled amounts of effluent from the reactor 171 to thereactors 151-1 through 151-n.

In a commercial-size plug flow reactor, the proportion of fresh feed andreaction products and byproducts varies continuously along the length ofthe catalyst bed. At the inlet there is 100% fresh reactant feed andzero reaction products and byproducts. As the fresh feed is consumed inthe catalyst bed of the reactor, the proportion of fresh feed decreasesand the proportion of reactant products and byproducts increaseslongitudinally along the catalyst bed. The multiple parallel-serialreactor arrangement of FIG. 7 can be used to perform a number ofdifferent kinds of experiments. For instance, all of the reactor stagescan contain the same catalyst and the composition of the feed can bevaried from stage to stage. Alternatively, the composition size orconfiguration of the catalyst particles can be varied from reactor stageto reactor stage in each of the reactor stages can receive the samefeed.

In accordance with the method of the invention, the longitudinalgradients in kinetics, mass transfer and heat transfer characteristicsfor the various reactions occurring within the catalyst beds, orcatalyst bed segments, for each of the upgrading steps forhydroprocessing dirty feedstocks are investigated with the use ofcomposite multistage series-connected fixed bed reactors thateffectively permit the segmenting of each individual catalyst bed, orbed segment, into successive longitudinal slices to permit the taking ofmeasurements to investigate the kinetic, mass transfer and heat transfercharacteristics for the different chemistries occurring within each ofsuch slices of the catalyst beds or bed segments. In the initial step ofremoving the hetero oxygen, sulfur and nitrogen atoms from theheterocyclic molecules, the dirty feedstocks and hydrogen are fed to theinlet of the first reactor stage of a first composite multistageseries-connected plug flow reactor in which the catalyst beds contain acatalyst, typically a metal sulfide, that is capable of catalyzing thereaction of hydrogen with the dirty feedstocks to remove the heteroatomsfrom the heterocyclic molecules in the feedstock. The catalyticchemistry performed in such first composite multistage series-connectedreactor initially forms alcohols, thiols and amines as intermediatesthat are liquid at the reaction conditions. These intermediates are thenconverted to H₂O, H₂S, and NH₃ that are highly volatile at reactiontemperatures and removed from the system as gases. If the reactionconditions are such that the intermediates were not fully converted,they would remain in the system and cause problems in the followingprocessing steps by poisoning the catalysts. By segmenting theheterocyclic molecule processing catalyst bed into a series of three ormore longitudinal catalyst bed slices in separate series-connectedreactor stages, sampling of the effluents of each of the reactor stagesand measuring the amounts of reactants and intermediate and finalproducts in the effluents of each of the stages in order to determinethe extent to which the hetero-atoms have been removed from theheterocyclic molecules and intermediate and final products andbyproducts of the reaction have been formed in each reactor stage, onecan optimize the processing conditions for the step to ensure, e.g.,that such intermediates do not remain in the system, but are convertedto the volatile products. The compositions of the effluents can bemeasured by gas or high-pressure liquid chromatic measurements combinedwith mass spectroscopy. Alternatively, the presence of sulfur can bemeasured using pulsed flame photometric detectors. Oxygen it can bemeasured using GCMS or O¹⁷NMR, and nitrogen can be measured usingchemiluminescence techniques.

Similarly, the following hydroprocessing steps are each performed insuccessive composite multistage series-connected fixed bed reactors,each containing three or more reactor stages. Each of such followingsuccessive hydroprocessing steps involves feeding effluent from the laststage of the preceding composite multistage reactor and hydrogen to thefirst stage of the following composite multistage reactor, and measuringthe concentrations of the feed to said following composite multistagereactor and the reaction products and byproducts in the effluents ofeach of the reactor stages of said following composite multistagereactor. The segmentation of the catalyst beds of each of the steps usefor hydroprocessing dirty feedstocks and the sampling and measuring theconcentration of the various components in the effluents of each of thereactor stages allows one to develop a model for optimizing theprocessing conditions for each of the steps using small-scale equipmentsuch as described above, which model is applicable to the analysis ofthe performance of commercial scale catalytic hydroprocessing systemsuseful for hydroprocessing dirty feedstocks.

In the step of using hydrogen to saturate the polynuclear aromatics, theprogress of the catalytic reaction in each of the reactor stages can beinvestigated by sampling the effluents of each of the reactor stages andtaking measurements using high-resolution P-NMR or C¹³NMR. The progressof the step of cleaving the carbon-carbon bonds, and the isomerizationstep, can be measured using gas chromatography with high-resolution massspectroscopy or high-resolution C¹³NMR measurements on the effluents ofthe relevant reactor stages. The identity and concentration of varioushydrocarbons can be measured using GCMS.

Among the process variables that can be investigated using the method ofthe invention is the effect of low levels of heteroatoms on thefunctionality of the catalysts in the process stages following theinitial heterocyclic molecule treatment step. This can be performedusing, e.g., a probe reactor in a manner described above with relationto FIGS. 2 through 4 or 6 of the drawings.

The data gathered based on these measurements allows the development ofpredictive models using small-scale reactors that are able to describebehavior applicable to large-scale catalytic hydroprocessing systems.

Kinetics

Heretofore, it has been the practice to measure the kinetics of a plugflow catalytic system only by measurements taken at the inlet and theoutlet of the catalyst bed, so that the measurements are averaged overthe length of a catalyst bed. In analyzing the kinetic performance ofsuch a reactor it was necessary to make assumptions concerning thekinetic order of the reaction. Typically, it was assumed that the orderof the reaction remained constant along the length of the catalyst bedin the reactor. Applicants have found that this assumption was in manycases incorrect. With the use of the multistage series-connected plugflow reactor of the present invention as described above with relationto any of the FIGS. 1 through 7, it is possible to investigatelongitudinal variations in the kinetics of a plug flow catalytic systemalong the length of the composite catalyst bed of the reactor.

Using the multiple parallel-serial reactor arrangement illustrated inFIG. 7 of the drawings as an example, the multistage series-connectedreactor of the present invention can be used in accordance with themethod of the invention to develop scale-up data for investigating theintegral, differential and intrinsic kinetics of a plug flow catalyticreactor system as a function of the longitudinal position along thecatalyst bed of the reactor. For example, to determine the integralkinetics of a fixed bed reactor system, the catalyst beds in the reactorstages of modules 151, 155 and 157 and the reactor 171 can contain thecatalyst intended for use with the system. The parallel reactor stages151-1 through 151-n in the module 151 receive varying proportions offresh feed from the source 167 and effluent from the reactor 171. Forinstance, the valves 169-1 through 169-n and valves at 173-1 through173-n can be set such that reactor stage 151-1 receives 100% fresh feedand no effluent, and the reactor stages 151-2 through 151-n receivesuccessively decreasing proportions of fresh feed and increasingproportions of effluent. In this arrangement, the successive reactorstages 151-1 through 151-n are equivalent to successive,longitudinally-spaced slices of the catalyst bed of a fixed bed reactor,with reactor stage 151-1 being equivalent to the slice at the inlet ofthe catalyst bed and reactor stages 151-2 through 151-n operating atconditions equivalent to slices of the catalyst bed positioned atsuccessive longitudinal positions along the composite bed. The reactorstages in modules 155 and 157 can be used to provide data for slices ofthe catalyst bed being scaled-up that are intermediate the slices of thesuccessive reactor stages in module 151. For example, if reactor 171 isoperated at 90% conversion, its effluent will contain 10% of the amountof fresh feed at its inlet with the remainder of the effluent beingreaction products and byproducts. If reactor stage 151-2 receives 88%fresh feed and 12% effluent from the reactor 171, the composition of thefeed at the inlet to reactor stage 151-2 will be 89.2% fresh feed withthe remainder being reaction products and byproducts. If the reactorstages 151-1, 155-1 and 157-1 are each run at 3% conversion, theireffluents will contain 97% fresh feed, 94.1% fresh feed and 92.3% freshfeed, respectively, with the remainder being reaction products andbyproducts. Thus, the compositions and proportions of fresh feed andreaction products and byproducts in the reactor stages in modules 151155 and 157 are equivalent to those at successive longitudinal slices inthe catalyst bed of a fixed bed reactor.

In order to determine the integral kinetics of the catalytic system as afunction of longitudinal positions in the catalyst bed, it is necessaryto analyze the inlet feed stream and composition and outlet feed streamand composition, normalized, for instance to STP per standard leader offeed, at each of the successive longitudinal slices of the catalyst bed.In this manner it is possible to determine the kinetics of the reactionat each effective slice of the full commercial scale system. If theresulting rate constant vs distance from inlet curve is a straight line,the integral kinetics of the system is a constant along the length ofthe catalyst bed. If the slope of the line changes from slice to slice,this would be indicative of a change in kinetics (and likely catalyststate) as a function fo distance from the inlet.

If the resulting kinetic curve on the log-log plot is not a straightline, then the integral kinetics of the system varies along the lengthof the reactor catalyst bed. In this case, it is necessary to do aregression analysis to fit the curve to an equation relating thereaction rate to the concentration of feed. Differentiating thatequation, either graphically or mathematically, gives the Rate ModelCorrelation as a function of longitudinal position along the catalystbed. A representative graphic technique is discussed in GraphicalMethods for Data Analysis, John M. Chambers, Chapman and Hall, May 1983,ISBN: 0412052717.

In order to determine the effects of temperature and pressure on theintegral kinetics of the system, the above-described experiment can berun at different temperatures and at different pressures. The experimentcan also be run using different size catalysts. For example, theexperiment can be run using the intended commercial size and shapecatalyst and also with a diluted crushed or powdered catalyst.

The intrinsic and differential kinetics, free of mass transfer and heattransfer effects, of the composite multistage series-connected fixed bedcatalytic system of the invention can also be investigated for purposesof scale-up to a commercial system using the systems depicted in FIGS.1-7 of the drawings. Using the system depicted in FIG. 7 as an example,the catalyst beds of the reactor stages include very finely crushed orpowdered catalyst particles in order to avoid mass transfer effects, andthe catalyst is highly diluted to avoid heat transfer effects.Additionally, the diameter of the reactor should preferably be small,typically about 5 to 12 millimeters to further avoid heat transfereffects. This can be accomplished by using a smaller diameter reactor orby using a heat conductive sleeve in each reactor stage to reduce itsdiameter. The depth of the catalyst bed in each of the reactor stages istypically between about 5 and 10 centimeters. The same series ofmeasurements and calculations are performed as described above fordetermining the integral kinetics of the system. In determining thedifferential kinetics of the system the amount of conversion in eachreactor stage should be very small, e.g. less than 20 percent,preferably about 2 to 5 percent. The measurements can be performed atdifferent temperatures and pressures in order to investigate the effectsof temperature and pressure on the intrinsic and differential kineticsof the system.

While these kinetics measurements have been described with relation toFIG. 7, it would also be possible to use the other disclosed reactorsystems such as that described with relation to FIG. 1 or 5 of thedrawings, using enough series-connected reactor stages to give thenecessary of longitudinal information along the composite catalyst bed.A significant advantage of the system of FIG. 7 of the drawings is thatthe use of the reactor 171 to supply the effluent to all of the reactorstages in module 151 means that each of the reactor stages in the module151 receives exactly the same reaction products and byproducts and traceelements, thereby replicating actual reactor conditions more exactly andeliminating errors resulting from variations in the composition of thefeed to the reactor stages. Additionally, the composition of the inputsand outputs from all of the reactor stages can be sampled substantiallysimultaneously to give a snapshot of the reactor's performance at agiven moment. The sampling of the composition of the inputs and outputsfrom the reactor stages can also be repeated periodically while thereactor system continues to operate thereby investigating theperformance of the reactor system as a function of time on stream to seewhat aspects of the reactor performance change and in what longitudinalzones of the overall catalyst bed the changes occur. This data is usefulin investigating the catalyst stability, among other things.

The kinetics of deep desulfurization is governed by the extent to whichdesulfurization (HDS) occurs by direct sulfur extraction, or byhydrogenation of the sulfur-containing molecule followed by sulfurextraction. The direct route is primarily inhibited by hydrogen sulfide,and the hydrogenation route by specific nitrogen-containing compounds.Certain catalysts are known to exhibit different preferences for the tworoutes, and this is ultimately important to the overall amount ofhydrogen that will be required to process a given amount of sulfurcontaining feed. Thus, a detailed understanding of the kinetics of deepdesulfurization is used to select the most suitable catalyst for a givenservice and to evaluate the relative advantages of a revamp versus agrassroots unit with respect to investments, hydrogen costs and productproperties.

The problem of deep removal of sulfur has become more serious due to thelower and lower limit of sulfur content in finished gasoline and dieselfuel products by regulatory specifications, and the higher and highersulfurcontents in the crude oils. A survey of the data on crude oilsulfur content and API gravity for the past two decades reveals a trendthat U.S. refining crude slates continue towards higher sulfur contentsand heavier feeds. Consequently there is a growing need for newer andmore efficient catalysts for deep desulfurization, and the presentinvention enables such materials to be developed in a more timely, costeffective and efficient way.

Mass Transfer

Methods of investigating the mass transfer characteristics of acatalytic process in a plug flow reactor, such as a fixed bed reactor,typically involve a comparing the conversion versus residence timecharacteristics at a given set of operating conditions of a finelycrushed with that of a commercial-size catalyst. The crushed catalyst isscreened to a narrow particle size range, preferably one that is closeto the minimum obtainable catalyst particle size that still retains itscatalytic properties. This minimum catalyst particle size depends on thecharacteristics of the specific catalyst being used, and can bedetermined by simple experimentation. In the more simple method fordetermining the mass transfer characteristics, the finely crushed andscreened catalyst is assumed not to have any mass transfer limitations,so that any difference in the conversion versus residence timecharacteristics between the crushed catalyst and the commercial-sizecatalyst is assumed to be the result of mass transfer limitations. For agiven feed, the effluent of the two reactors is sampled to determine theamount of conversion. Alternatively, the input flow rates of the tworeactors can be adjusted (i.e., the input flow rate to the crushedcatalyst in reactor is increased, or the input flow rate to thecommercial-size catalyst reactor is decreased) so that each of thereactors has the same percentage conversion, and that difference inresidence times is attributed to mass transfer limitations in thecommercial-size catalyst.

In a more rigorous and technically exact method of determining the masstransfer characteristics of a commercial-size catalyst, the finelycrushed catalyst is not assumed to have zero mass transfer limitations,and the Thiele Modulus of the commercial catalyst is determined from theratio of the observed reaction rates of the crushed and commercial-sizecatalysts and the ratio of their particle sizes. The EffectivenessFactor for the commercial-size catalyst can then be determined from aplot of the effectiveness factor versus the Thiele Modulus. This methodis described in Hougen and Watson, Chemical Process Principles, PartIII, Kinetics and Catalysts, pp. 998-1000, Wiley, March 1966, which isincorporated herein by reference.

The intra-particle diffusion effect on hydrodesulphurization of modelcompounds such as dibenzothiophenic compounds in light cycle oil can bedetermined experimentally by varying the average particle size anddetermining the effect of that change on overall catalyst activity andselectivity. The catalyst effectiveness factors can then be determinedfor a full-size catalyst and its crushed catalyst particles undersimilar operating conditions. In the full-size commercial catalyst,significant intra-particle diffusion resistance is possible, resultingin reduced sulphur removal reaction rates for all the dibenzothiophenicand related sulfur containing reactants in the feed. The effectivenessfactors are a good measure of the degree of intraparticle mass transferresistance and need to fall in the range of 0.3 or higher, preferably0.8 to 0.9 to achieve desired performance. It is also important todetermine the effect of intraparticle mass transfer resistance as afunction of the geometric properties of sulfur bearing reactants andthis can be achieved by use of model DBT compounds with varioussubstituent groups that increase their molecular size. A problem withboth of these methods is that they does not give any informationconcerning longitudinal variations in mass transfer performance alongthe reactor catalyst bed and basically assumes that the mass transfercharacteristics are uniform from input to output. This assumption isincorrect for many catalytic systems, and the inability to investigatethe longitudinal variations in mass transfer characteristics in a fixedcatalyst bed has meant that information which would allow theoptimization of the catalyst bed along its length has not beenavailable.

In accordance with the present invention, the catalyst beds of the plugflow reactors are segmented longitudinally into at least threeseries-connected stages and the effluent of each of the stages issampled to determine the amount of conversion occurring in eachlongitudinal segment of the catalyst bed. Referring again to FIG. 1 ofthe drawings, in accordance with the present invention, each of thereactors 11 and 35 includes three or more reactor stages with samplingvalves between the output of each stage and the input of the succeedingstage for measuring the content of the effluent of each stage. Thetemperature control device 33 maintains both of the reactors 11 and 35in a common thermal environment. The reactors 11 and 35 both receive theidentical reactant input feed from the source 31. In performing a basicmass transfer investigation, the sources 55, 57 and 59 are preferablynot used. The catalyst beds 19, 21 and 23 in reactor stages 13, 15 and17 of reactor 11 contain a finely crushed and screened or powderedcatalyst mixed with enough inert diluent particles so that the operationof the reactor 11 is essentially isothermal

The catalyst beds 43, 45 and 47 in reactor stages 37, 39 and 41 ofreactor 35 are composed of commercial-size catalyst particles that arealso mixed with a lesser percentage of inert diluent particles so thatthe operation of reactor 35 is also essentially isothermal. Toinvestigate the longitudinally-dependent mass transfer characteristicsof the commercial-size catalyst in accordance with the method of theinvention, each of the reactors 11 and 35 receive the identical reactantfeed from the source 31 and the pressure and the feed rate for each ofthe two reactors is held constant. The conversion versus residence timerelationship is obtained for each stage of the reactors 11 and 35 fromthe difference in the amount of reactant feed at the inlet and outlet ofeach reactor stage and the flow rate, for a given set of operatingconditions.

In the simplified method of determining mass transfer limitations, theEffectiveness Factor for the commercial-size catalyst is obtained forthe commercial-size catalyst at each stage of the reactor 35 by takingthe ratio of the Observed Reaction Rates of the commercial-size catalystand the crushed catalyst for each reactor stage. The Observed ReactionRate is obtained for each reactor 11 and 35 by plotting the cumulativeconversion of reactant and corresponding cumulative appearance of theproduct and byproducts (if any) versus residence times at the outputs ofthe reactor stages of each reactor and fitting curves to the data usingwell-known techniques. See, e.g., Graphical Methods for Data Analysis,John M. Chambers, Chapman and Hall, May 1983, ISBN: 0412052717. Theslope of the resulting curve for the product at any residence time orconversion level for one of the reactors 11 or 35 is the ObservedReaction Rate, K_(o) (conversion per unit of residence time) for suchreactor for such product. If mass transfer were not limiting, the K_(o)would be independent of particle diameter. A comparison of the plots ofK_(o) versus conversion for the two reactors defines the longitudinalareas of the composite catalyst bed of the reactor 35 containing thecommercial-size catalyst in which mass transfer through the catalystpores is limiting. The Effectiveness Factor for a catalyst in a reactoris equal to the K_(o) divided by the Intrinsic Reaction Rate, for suchcatalyst in the reactor. In the simplified method, the crushed catalystis assumed not to have any mass transfer limitations, so that its K_(o)is equal to the K_(i) for the catalyst. Therefore, the EffectivenessFactor for the commercial-size catalyst at any point along the compositecatalyst bed of reactor 35 is equal to the ratio of the K_(o) of thecommercial-size catalyst to that of the crushed catalyst at such pointalong the catalyst beds.

If the Hougen and Watson method is used, the K_(o) of the crushedcatalyst is not assumed to be equal to the K_(i). According to thismethod, it is possible, using the graph of FIG. 8 of the drawings, todetermine the Thiele Modulus for the commercial-size catalyst at anypoint along the catalyst bed from the ratio of K_(o)'s at such point andthe ratio of the particle diameters of the commercial-size and crushedcatalysts. For instance, if the ratio of the particle diameter of thecrushed catalyst to that of the commercial-size catalyst is 0.2, and theratio of K_(o) of the commercial-size catalyst to that of the crushedcatalyst is 0.34 at a given point along the catalyst beds, the ThieleModulus at that point is about 9. Using the graph of FIG. 9, theEffectiveness Factor for the commercial-size catalyst at that pointalong the composite catalyst bed of reactor 35 is about 0.27. Thedetermination of the longitudinally dependent Effectiveness Factor forthe catalyst bed containing the commercial-size particles can beperformed repeatedly during running of the reactors 11 and 35 todetermine the effect of time on stream on the mass transfercharacteristics of the fixed bed catalyst system. The measurements canalso be repeated at different operating conditions of temperature andpressure in order to investigate the longitudinally dependent effects ofchanges in these parameters on the mass transfer characteristics of thecomposite catalyst bed of the plug flow reactor 35.

Because the Effectiveness Factor is the ratio of K_(o) to the it ispossible to calculate the K_(i) for a catalyst from the EffectivenessFactor and the K_(o) for a given longitudinal point along the catalystbed. Since K_(i) is the same for the crushed and commercial-sizecatalysts, the Effectiveness Factor for the commercial-scale catalyst atany point along the catalyst bed can be determined from the K_(o) forthe crushed catalyst at that point and the K_(i).

For reactions in which different reaction pathways are possible indifferent longitudinal portions of the catalyst bed of the plug flowreactor, e.g., conversion of sulfur or nitrogen containing feedstocks,carburization, or the production of methane via hydrogenolysis, it isimportant also to characterize the behavior of the different kineticpathways producing the product and various byproducts that can exist forthe system as they vary along the length of the composite catalyst bedof the reactor in order to explore the longitudinally dependent kineticand mass transfer space for the system, and to distinguish between theoccurrence of mass transfer and kinetic effects in the system. When thisspace has been explored, the mass transfer performance of reactant toproduct for the system operating at a given set of conditions thatinvolve an optimal set of trade-offs for the particular catalyst can beinvestigated.

An example of the opportunity to optimize the longitudinalcharacteristics of a catalyst bed of a fixed bed reactor afforded as aresult of the data obtained by the method of the present invention isillustrated in connection with the graph in FIG. 10 of the drawings.This graph depicts what is believed to be a typical relationship betweenthe Effectiveness Factors and conversion rates for crushed andcommercial-size catalysts in a fixed bed reactor working at a given setof operating conditions of temperature and pressure and with a commonreactant feed. Mass transfer limitations are clearly present up to thepoint in each of the fixed bed reactor catalyst beds at which about 50to 60% conversion has occurred, but are not present at the portions ofthe catalyst beds at which greater than about 70% conversion hasoccurred. The greater mass transfer limitations, evidenced by of thelower Effectiveness Factor, of the bed containing the commercial-sizecatalyst particles is believed to reflect the differences in the lengthsof the reaction pathways in the crushed and commercial-size catalystparticles. This suggests that buildup of material, such as wax, in thecatalyst pores is present at the portion of the reactor catalyst bed atwhich lower conversion has occurred, i.e., close to the inlet of thereactor bed where the catalyst experiences almost entirely fresh feed,but not present at lower portions of the catalyst bed at which higherconversions have occurred.

In a hydroprocessing reactor, a lower Effectiveness Factor may result inan undesirable higher methane make and/or carbon overlayer formation.Thus, particularly in the reactor containing the commercial-sizecatalyst bed, the upper portions of the catalyst bed would be producingsubstantial amounts of methane. This results in a much lower diffusivityof the reactant gases in such pores, so that the active sites within thecatalyst become starved of reactants and begin generating large amountsof methane. In order to optimize the catalyst bed structure of the fixedbed reactor to avoid the undesirable high methane make in the inletportions of the catalyst bed, it is possible, for instance, to use aless active catalyst in that portion of the reactor bed, which wouldgenerate lesser amounts of methane.

As an alternative to using crushed in commercial size catalyst particlesof different sizes in an investigating the mass transfer characteristicsof the catalyst bed in a plug flow reactor, is possible to use the samesize particles with different levels of catalyst loading. The particleswould be made up of finely crushed or powdered catalyst dispersed andinert diluent such as alumina or silica. The powder or finely crushedcatalyst is uniformly mixed with the finely crushed inert diluent,formed into particles of a given size and sintered. Particles in whichthe catalyst concentration is selected to be relatively low cancorrespond to the crushed catalyst in the method described above.Particles in which the catalyst concentration is relatively higher cancorrespond to the commercial-size catalyst. The concentration ofcatalyst within the particles appropriate for the particles tocorrespond to crushed catalyst or commercial-size catalyst depends onthe activity of the catalyst and the nature of the reaction.

In scaling-up a reactor to commercial size, is preferable to confirm themass transfer characteristics determined under isothermal conditions inthe manner described above in an adiabatic reactor. In an adiabaticreactor, the amount of diluent for the commercial-size catalyst isreduced and the tube diameter is controlled so that its thermalperformance mirrors that expected for the commercial-size reactor.

In investigating mass transfer effects in a hydroprocessing reactor, asan alternative to plotting the reaction rate versus conversion orresidence times, is to plot the methane selectivity versus conversion.Methane selectivity is greater when mass transfer limitations exist.Mass transfer is an issue only in those parts of the reactor wheremethane selectivity is widely different for the commercial and crushedcatalyst. Between about 35% and 80% conversion, the methane selectivityis very low. In this region, mass transfer is not an issue. In theportion of the catalyst bed where above about 80% conversion hasoccurred, the methane selectivity increases rapidly and the reactionrate slows down for both the crushed and the commercial catalyst. Thisis an indication that something other than mass transfer effects islimiting the catalyst activity and increasing the methane selectivity.

Heat Transfer Effects

Understanding the heat transfer performance of a plug flow reactor iscritical to maximizing the productivity at which the reactor can be run.The temperatures in the catalyst bed of a plug flow reactor can varyboth longitudinally and laterally within the catalyst bed. Forendothermic reactions, there can be the need to get heat into cold spotsin the catalyst bed or the reaction may shut down.

The reactor system illustrated in FIG. 2 of the drawings can also beused to investigate heat transfer characteristics of a plug flow reactorsystem. For example, the catalyst beds in the reactor stages 13, 15 and17 of the reactor 11 can contain a mixture of crushed catalyst and inertdiluent particles, and the catalyst beds in stages 37, 39 and 41 of themultistage reactor 35 can contain mixtures of full-size catalystparticles and inert diluent particles. And both cases the ratios ofcatalyst particles to inert diluent particles are selected so that thereactor's 11 and 35 operate substantially isothermally. The catalystbeds of the reactors 11 and 35 are instrumented with thermocouples (notshown) to measure in the temperatures at successive longitudinalpositions along the catalyst beds, both in the central portion of thebed cross-section and near its periphery. In addition, the effluent ofeach of the reactor stages is sampled by sampling valves 25, 27 and 29of multistage reactor 11 and sampling valves 49, 51 and 53 of multistagereactor 35. Lateral heat transfer effects can be further studied byinserting conductive sleeves in the reactor stages in order to decreasethe catalyst bed diameter so that the heat generated in the centralportion of the bed has less distance to travel to the heat sink formedby the reactor walls and the temperature control device 33 surroundingthe reactor walls. Successively thinner heat conductive sleeves can beused to incrementally increase of the diameter of the catalyst bed untilthe bed diameter is such that the heat that cannot be adequately removedfrom the central portion of the bed through the reactor walls.

Temperature and product measurements are preferably a repeated fordifferent reactor flow rates, pressures, and productivities, both atStart of Run and during the reactor's time on stream as the reactorlines out. The effect on heat transfer characteristics and other processparameters, such as conversion, selectivity and kinetics, of usingcatalyst particles of various sizes and shapes in the catalyst bed canalso be investigated using the method of the invention. The dataobtained from such measurements permits one to investigate and gain anunderstanding of how the heat transfer properties of the reactor systemaffect reactor performance over the entire multivariable space in whichthe commercial-size reactor might operate.

Referring now to FIG. 11 of the drawings, there is illustrated inalternative embodiment of the apparatus of the invention which can beused for investigating the longitudinally dependent mass transfer,kinetics and heat transfer characteristics of a fixed bed reactor. Theplug flow reactor 201, in this case a fixed bed reactor, contains a bed203 of commercial sized catalyst particles. Reactor 201 is supplied withfresh reactant feed from the source 205. Effluent from the reactor 201is supplied to fixed bed reactor stages 207-1 through 207-n throughcontrol valves 209-1 through 209-n for feeding controlled amounts ofeffluent from reactor 201 to such reactors. Each of the reactor is 207-1through 207-n contains a narrow catalyst bed 211-1 through 211-n ofcatalyst particles mixed with enough inert diluent particles so that thecatalyst beds operate in a substantially isothermal mode. The source 205also supplies controlled amounts of fresh reactant feed to the inlets ofthe reactor stages 207-1 through 207-n through control valves and 211-1through 211-n. The effluents from the reactor or stages 207-1 through207-n can be sampled by means of sampling valves 215-1 through 215-n.

If the reactor 201 is operated at a given conversion level, e.g. 80%,the input to the individual reactor stages 207-1 through 207-n canrepresent any degree of conversion from zero to 80% by using the controlvalves 209-1 through 209-n and 213-1 through 213-n to adjust the ratioof reactor 201 effluent to fresh feed being supplied to the individualreactor stages 207-1 through 207-n. Thus, if the valves 209-1 and 213-1are adjusted such that reactor stage 207-1 receives only effluent fromthe reactor 201, and the thickness of the catalyst bed 211-1 is suchthat it performs an additional 5% conversion on such effluent, thecatalyst bed 211-1 is equivalent to a cross-sectional slice of a fixedbed reactor in which the conversion between 80 and 85% takes place.Similarly, if the valves 209-2 and 213-2 are adjusted such that theinput to reactor stage 207-2 is equivalent to the effluent of a reactoroperating at 40% conversion, and the thickness of the catalyst bed 211-2is such that it performs an additional 5% conversion on such effluent,the catalyst bed to an 11-2 is equivalent to a cross-sectional slice ofa catalyst bed in which the conversion between 40 and 45% takes place.Thus, the catalyst beds 211-1 through 211-n can replicate theperformance of a cross-sectional slice of a fixed bed reactor positionedat any longitudinal position along the catalyst bed.

The catalyst beds 211-1 through 211-n need not all have the samecomposition. For instance, the beds 211-1 and 211-2 could containcrushed and commercial-size catalyst particles, respectively, in eachcase mixed with an amount of inert diluent particles such that the bedsoperate in isothermal mode. In this case the mass transfer, heattransfer and kinetics characteristics of a cross-sectional slice of acatalyst bed located at any longitudinal position in the catalyst bedcan be investigated. In a different application, the catalyst beds 211-1through 211-n could contain catalyst particles of different chemical orphysical composition. In order to prevent heat loss or gain in theeffluent from the reactor 201 being fed to the reactor stages 207-1through 207-n, the connecting tubing and valves are preferablysurrounded by insulating material and the entire system comprising thereactor 201 and the reactor stages 207-1 through 207-n can be surroundedby a temperature control device, or alternatively, the reactor 201 andreactor stages 207-1 through 207-n can be surrounded by separatetemperature control devices, depending on the needs of the application.Additionally, the reactant feed from the source 205 being supplied tothe reactor stages 207-1 through 207-n can be heated before it issupplied to such reactor stages by well-known indirect heating meanssuch as a coil in a sand bath or an infrared furnace (not shown) inorder to have the appropriate temperature conditions in the catalyst bedinlet portions of such reactor stages.

The apparatus disclosed in FIGS. 2, 4, 7 and 11 can also be used toinvestigate other operating parameters of a plug flow reactor forscale-up or other purposes in accordance with the method of theinvention. For example, the longitudinally dependent activitymaintenance of a catalyst bed can be investigated as a function of timeon stream under different conditions of temperature, pressure andcatalyst shape and size. Other longitudinally dependent processparameters that can be investigated using the method of the inventioninclude the effects of different space velocities, reaction products andby-products, different operating temperatures and pressures, time onstream, and different catalyst sizes and shapes, on matters such ase.g., conversion, productivity, kinetics and selectivity, and on changesin catalyst physical and chemical properties such as active site crystalsize, oxidation, and growth of an over-layer of support on the surfaceof the catalyst active sites.

Using present invention, the time for scale-up of the catalytic processfrom discovery to commercial scale application can be significantlyreduced. For example, in one particularly advantageous configuration,four multi-stage reactors of the type described above can be operated inparallel. In this embodiment, the stages of one of the reactors areloaded with crushed catalyst. This reactor provides Intrinsic ReactionRate and selectivity data. The stages of the second reactor are loadedwith commercial-size catalyst. The data from this second reactor can beused to define the degree of mass transfer limitation (EffectivenessFactor) based on a direct comparison of the relative residence times inthe reactors containing the crushed catalyst in the commercial-sizecatalyst required to achieve a given amount of conversion. By obtainingconversion data at a series of residence times, it is possible todetermine the Effectiveness Factor and hence the Effective Diffusivitywith conversion or residence time. This data also provides informationon the impact of mass transfer on selectivity. A third, probe reactorcan be operated in parallel with the previous two reactors. This probereactor can either be a shallow fixed bed reactor or a back-mixedreactor. Flow can be directed to the appropriate actor from any of thereactor beds in the previous two reactors. In addition, additional gasesor liquids can be added to the probe reactor to determine the rates ofadsorption or surface property changes on the catalyst. This informationcan provide valuable insight in modeling the fixed bed reactor. Finally,an adiabatic reactor can be operated in parallel to test the reactormodel developed from the previous reactors. Operation of the seriesreactors in this parallel mode allows for much faster generation of therequired scale-up data. In fact, all the required scale-up data,including deactivation and regeneration data, at one temperature can beobtained in one to two years, for a savings of several years ofdevelopment time. A further improvement to the experimental design wouldbe to operate several four reactor sets at the same time. These sets canbe operated at different temperature, pressure, and feed compositions.The set producing the optimum economics can be used for the commercialdesign. The cost of operating several parallel sets of series reactorssimultaneously is a small expense when compared to the potential savingsassociated with accelerating the scale-up of a new catalyst to afull-scale commercial operation. If the new catalyst results in a$1/barrel savings, a 100 thousand barrel/day plant will produce asavings of over $30 million per year. These savings would easily farmore than offset the cost of operating the parallel sets of seriesreactors.

In an adiabatic reactor, it is possible to produce hot spots in thereactor, which may cause the adiabatic reactor to run away. Also, in anadiabatic reactor, because reaction parameters, such as temperature,kinetics parameters, etc., can change continuously, it is difficult tomeasure the reaction parameters by direct measurement. Dividing anadiabatic reactor into multistage series-connected reactor stages canhelp determine reaction parameters at different locations along a flowdirection of the reactor, but it is difficult to keep continuities ofthe reaction parameters, especially temperature, between adjacentreactor stages.

Therefore, it is difficult to directly measure reaction parameters in anadiabatic reactor, and to exactly and securely determine reactioncharacteristics in the adiabatic reactor, such as kinetics, masstransfer, heat transfer etc.

FIG. 12 illustrates a schematic diagram of a composite multistagelaboratory scale plug flow reactor 607. The reactor 607 includes first,second and third series-connected reactor stages 61, 63 and 65, eachhaving a catalyst bed 62, 64 and 66. The reactor 607 further includes afresh reactant conduit 70 which connects an inlet of the first reactorstage 61 to a source 60, so that the source 60 can provide feeds, whichare normally fresh reactants, to the first reactor stage 61. The reactor607 further includes connecting conduits 71 and 72 to connect the firstand second reactor stages 61 and 63, and the second and the thirdreactor stages 63 and 65, respectively. A first sampling valve 67 isdisposed between the first and second reactor stages 61 and 63, and hasan output 601 to facilitate sampling effluents from the first reactorstage 61. Here in this document, a device is said to be disposed betweentwo stages of the reactor does not necessarily mean that the device isphysically disposed between the two stages of the rector but that thedevice is between the two stages of the reactor along a flow ofreactants. A second sampling valve 68 is disposed on the conduit 72 andhas an output 602 for sampling effluents from the second reactor stage63. A third sampling valve 69 is disposed between an outlet of the thirdreactor stage 65 and a device, such as a fourth reactor stage or aproduct accumulator (not shown) and has an output 603 for samplingeffluents from the third reactor stage 65. A sampling valve connected tothe fresh reactant conduit 70 may also be provided in order to permitanalysis of the feeds.

In one embodiment, the reactor stages 61, 63 and 65 are isothermalreactor stages, which are used together to simulate an adiabaticreactor. Thus, temperature control devices 604, 605 and 606 are providedto control the temperature of the reactor stages 61, 63 and 65respectively. A preheater (not shown) may be disposed between the source60 and the first reactor stage 61 to preheat the feeds from the source60 so that when the feeds flow into the first reactor stage 61, thefeeds have already reached a desired temperature for the feeds.Alternatively, the preheater can also be disposed in the first reactorstage 61.

In one embodiment, when using the isothermal reactor stages 61, 63 and65 to simulate the characteristics of an adiabatic reactor, thetemperature setting for each of the temperature control devices 604, 605and 606 should be determined first. Generally, for a given catalyticprocess, based on data derived from operating the adiabatic reactor inpractice, the temperature setting for the first temperature controldevice 604 and temperature variation in the first reactor stage 61 canbe determined. Then, based on information from the first reactor stage61, the temperature setting of the second temperature control device 605can also be determined, and so on. Thus, after the temperature settingsof each of the temperature control devices 604, 605 and 606 isdetermined, the reactor stages 61, 63 and 65 can be used to simulate thecharacteristics of the adiabatic reactor.

In this embodiment, the temperature of the temperature control devices604, 605 and 606 are defined as T1, T2 and T3, which are different fromeach other. Different catalytic processes may have different T1, T2 andT3 settings. Alternatively, a common temperature control device (notshown) can be provided to control the temperatures of reactor stages 61,63 and 65 together.

Thus, the isothermal reactor stages 61, 63 and 65 can respectivelysimulate successive catalyst bed slices of a catalyst bed of a largeradiabatic reactor. Thus, the characteristics of the catalyst bed, whichis simulated by the catalyst beds 62, 64 and 66, are determined. Becauseit is relatively easy to operate the isothermal reactor stages,characteristics associated with the larger adiabatic reactor can bedetermined by simulating the adiabatic reactor using the isothermalreactor stages. In this embodiment, the first, second and third reactorstages 61, 63 and 65 can be arranged upright.

For a particular catalytic process between at least two successivereactors, for example a particular catalytic process in a multistageseries-connected reactor stages, if an effluent fluid from one reactorstage is homogeneous, such as in a gas phase, transferring effluentfluid can be quite straightforward by using a properly sized and shapedtube connecting an outlet of one reactor stage to an inlet of afollowing reactor stage. In many catalytic processes, however, theeffluent from a reactor stage may be in a multiphase state, meaning thatit includes one or more gaseous fluids, which are fluids in gas phase(such as gases, vapors or mixtures of gases and vapors), and one or moreliquid fluids, which are fluids in one or more liquid phases (such aswater phase, oil phase, other immiscible phases and partial emulsionphases, etc.)

The multiphase fluid is often a multi-component fluid, each componentbeing in its own state, which can be a single-phase state or multiphasestate. If the multi-component fluid is in thermodynamic equilibrium, thefluid can be transferred directly by a tube connecting two successivereactor stages.

However, in certain catalytic processes, such as hydrodesulphurizationetc., the multi-component fluid may not be in thermodynamic equilibrium.So, when the multi-component fluid is transferred directly through thetube connecting the outlet of one reactor stage to the inlet of thefollowing reactor stage, the states of the components may vary duringthe transfer such that continuity or consistency of the fluid betweenadjacent two reactor stages may be broken. Thus, it is difficult to usethe multistage series-connected reactor stages to model a plug reactorand to measure and optimize the corresponding catalytic processes.

FIG. 13 illustrates a schematic diagram in accordance with oneembodiment of the present invention. In this embodiment, a catalyticprocess development apparatus includes a composite multistage laboratoryscale plug flow reactor 707 which includes first and secondseries-connected reactor stages 71 and 73. The reactor stages 71 and 73include catalyst beds 72 and 74, respectively. The catalytic processdevelopment apparatus further includes temperature control devices 701and 702 disposed on the reactor stages 71 and 73 respectively, and afresh reactant conduit 77. The fresh reactant conduit 77 is connected aninlet of the first reactor stage 71 to a source 70 so that the source 70can provide feeds which are normally fresh reactants to the firstreactor stage 71. In this embodiment, the catalytic process developmentapparatus further includes a separator 703, first and second effluentconduits 78, a gas conduit 75 and a liquid conduit 76. The first conduit78 is connected an outlet of the first reactor stage 71 to an inlet ofthe separator 703. The gas conduit 75 and the liquid conduit 76 connectthe separator 703 to an inlet of the second reactor stage 73. The secondeffluent conduit 78 connect an outlet of the second reactor to afollowing device (not shown), such as another separator. The reactantsfrom the source 70 are fed into the first reactor stage 71. A multiphaseeffluent fluid from the first reactor stage 71 is sent into theseparator 703, wherein gaseous fluid(s) in the multiphase fluid areseparated from liquid fluid(s), and both are introduced into the secondreactor stage 73 through the gas conduit 75 and the liquid conduit 76respectively.

Referring to FIG. 3, the catalytic process development apparatus furtherincludes a flow restrictor 705 disposed on the gas conduit 75 to controlflow resistance in the gas conduit 75, resulting in a gas pressuredifference (pressure drop) ΔP between two sides of the flow restrictor705. Assuming a gas pressure in the first reactor 71 and the separator703 is P1, a gas pressure in the second reactor 73 is P2. Thus, P1>P2due to the flow restrictor 705, and ΔP=P1−P2.

In one embodiment, ΔP is large enough so that it can drive the liquidfluid in the separator 703 to enter into the liquid conduit 76 and toflow into the second reactor stage 73 but is also small enough so thatit can not affect reactions in the second reactor stage 73. The flowrestrictor 705 can be a restricting valve, an orifice, or otherrestricting means etc. When properly sized and shaped, the gas conduit75 can function as the flow restrictor 705. The flow resistance of thegaseous fluid can be adjusted by many ways, such as electrical,electromagnetic, pneumatic, mechanical or thermal ways etc., which arefamiliar to those ordinary skills in the art. The electromagnetic waysare preferred.

Additionally, the catalytic process development apparatus furtherincludes a differential pressure sensor (not shown) disposed across theflow restrictor 705 or two ends of the gas conduit 75 to measure the ΔP.Combined ΔP and physical properties of the gaseous fluid, informationabout a mass flow rate of the gaseous fluid can be determined.

In one embodiment, if ΔP is too small, the liquid fluid can not flow butaccumulate in the separator 703. If ΔP is too large, the liquid fluidmay keep flowing until all the liquid fluid in the separator 703 istransported to the second reactor stage 73. When the liquid fluid in theseparator 703 is drawn out, the gaseous fluid may flow through theliquid conduit 76. Thus, ΔP is reduced due to an extra pathway for thegaseous fluid. Then, the liquid fluid begins to accumulate in theseparator 703 and blocks the liquid conduit 76. Subsequently, the ΔPrestores to a desired value little by little, and the liquid fluidstarts to flow again. Thus, the flow rates of the gaseous and liquidfluids may fluctuate with respect to time because of fluctuation of theΔP, which is disadvantageous to the second reactor stage.

In a preferred embodiment, the catalytic process development apparatusincludes a liquid level sensor 706 disposed in the separator 703. Theliquid lever sensor 706 monitors variation of a liquid level 704 in theseparator 703. Liquid sensor signals from the liquid level sensor 706are used to control the flow restrictor 705 to generate a suitable ΔP todrive the liquid fluid in such a manner that the liquid level 704 ismaintained at a desired substantially constant level. Thus, thefluctuation of the fluids in the separator 703 can be eliminated. Whenthe liquid fluid is transferred stably through the liquid conduit 76,the liquid mass flow rate information can also be obtained by using themeasured ΔP in combination with physical properties of the liquid fluid.

In one embodiment, in certain low pressure reactions including lowpressure FT synthesis etc., a small pressure drop ΔP may still be toobig to tolerate, especially when the reactor stage is long or there aremany reactor stages. Additionally, in the process of adjusting ΔP tomaintain the liquid level 704 by the liquid level sensor 706 and theflow restrictor 705, the fluctuation of ΔP may also affect liquid flowin the first reactor stage 71.

FIG. 14 illustrates a similar schematic diagram as the diagram of FIG.13. In this embodiment, the flow restrictor 705 is removed from the gasconduit 75, so, there is no pressure drop ΔP on the gaseous fluid.Meanwhile, a liquid pump 707 is disposed on the liquid conduit 76. Theliquid level signals are used to control the liquid pump 707 to maintainthe liquid level 704 at the desired constant level. Additionally,because an output pressure of the liquid pump 707 is approximately equalto its input pressure, it does not create a pressure drop between thefirst and the second reactor stages 71 and 73.

In this embodiment, the liquid pump 707 includes a positive displacementpump or a centrifuge pump etc. Additionally, the liquid pump 707 canhave metering capability, which can be used to obtain the liquid flowrate information directly. In order to cause the liquid fluid to bedistributed uniformly in the second reactor stage 73, a sprayer orsimilar spraying devices (not shown) can be adopted inside the reactorstage 73. Alternatively, a check valve (not shown) may be disposed onthe liquid conduit 76 and located behind the liquid pump 707 to preventthe liquid fluid in the liquid conduit 76 from reflux.

In the embodiments of the present invention, the gaseous fluid and theliquid fluid in the effluent of the first reactor stage 71 are separatedin the separator 703, and then transported to the second reactor stage73. Thus, possible interactions between the gaseous fluid and the liquidfluid in the effluent during transport can be minimized, and thepotential of altering the states of the components in the effluent byfluid distribution and recombination processes can be reduced. Thecontinuity or consistency of the components of the fluid can bemaintained between the first and second reactor stages 71 and 73.Additionally, separation of the gaseous fluid and the liquid fluid alsomakes it easy for sampling the fluids for species analysis, whethercontinuously or intermittently, on-line or off-line.

As mentioned above, in certain catalytic processes, there are differenttypes of liquid phases for the multiphase effluent fluid. In one exampleof the FT synthesis, its effluent may contain water phase liquid(s) andoil phase liquid(s). In order to transport such multiphase fluiduniformly, an agitation device (not shown) can be provided to causehomogenization of the multiphase fluid. The agitation device may includea mechanical stirring device, a magnetic stirring device or anultrasonic stirring device etc. In one embodiment, the ultrasonicstirring device is provided, which can be installed near a bottom of theseparator 703. The ultrasonic stirring device can provide sufficienthomogenization of the liquid fluid, while having minimum interference tothe performance of the liquid level sensor 706 and also withoutsignificantly increasing liquid temperature.

Referring to FIGS. 13-14, if the separator 703 is operated in atemperature which is higher than that of the first reactor stage 71,portions of volatile species in the liquid phase in the separator 703may be evaporated and enter into the gas phase so as to alter the statesof the species. If the separator 703 is operated in the temperaturewhich is lower than that of the first reactor stage 71, portions ofvapors in the gas phase in the separator 703 may be condensed and enterinto the liquid phase so as to also alter the states of the species. Asa result, variations in the effluent from the first reactor stage 71 canbe produced during its transfer to the second reactor stage 73.Therefore, for certain catalytic processes, it is preferred that thetemperature of the separator 703 is the same as that of the effluentfrom the first reactor stage 71. Thus, the states of the species of theeffluent are preserved.

Referring to FIG. 15, for example, in order to keep the temperature ofthe separator 703 being the same as that of the effluent of the firstreactor stage 71, the separator 703 is integrated into the first reactorstage 71. The integrated first reactor stage 71 and the separator 703can enjoy operation simplicity and also minimize the potential ofaltering the states of the components.

The composite multistage reactor 707 can include three or moreseries-connected reactor stages. The outlet of each of the reactorstages can connect to a separator. The separator and the reactor stagecan be separate from or integrated with each other. All the reactorstages can also be arranged upright along a vertical line.

1. A method for determining a set of operating parameters for acommercial scale plug flow catalytic process and reactor system forhydroprocessing dirty feedstocks, comprising the steps of: removingheteroatoms from heterocyclic molecules in said dirty feedstock byfeeding selected partial pressures of said feedstock and hydrogen to theinlet the first reactor stage of a first composite multi-stageseries-connected laboratory scale plug flow reactor including at leastthree reactor stages, the catalyst beds of each of said reactor stagesincluding catalyst particles capable of catalyzing the removal byhydrogen of heteroatoms from said heterocyclic molecules, saidheteroatom removal process being implemented at a selected set ofoperating conditions of temperature, pressure, and reactant and reactionproduct flow rates, and the catalysts in the catalyst beds of saidlaboratory scale reactor stages having selected sets of characteristics;sampling the effluents of each of said reactor stages; measuring theconcentration of heterocyclic molecules in said dirty feedstock in theconcentrations of heterocyclic molecules and intermediate and finalproducts and by products of the catalytic reaction in the effluents ofeach of said reactor stages; repeating steps (a) through (c) atdifferent selected sets of said operating conditions, and/or atdifferent selected sets of characteristics of the catalysts in thecatalyst beds of said laboratory scale reactor stages; and using theresults of said measurements obtained in one heteroatom removaloperation to influence the selection of catalyst bed characteristics andoperating parameters in a subsequent heteroatom removal operation forimproving the productivity and selectivity of the laboratory scaleplug-flow reactor.
 2. The method of claim 1 wherein said measuring stepincludes the step of measuring the amounts of thiols, amines andalcohols in the effluents of each of said reactor stages.
 3. The methodof claim 1 further including the steps of: saturating polynucleararomatic molecules in the effluent of the said first laboratory scalereactor by feeding selected partial pressures of effluent from the laststage of said first multistage laboratory scale reactor and hydrogen tothe inlet of the first reactor stage of a second multistageseries-connected laboratory scale plug flow reactor in which each of thereactor stages of said second reactor includes a catalyst bed containingcatalyst particles capable of catalyzing the saturation by hydrogen ofsuch polynuclear aromatic molecules, said polynuclear aromatic moleculesaturation process being implemented at a selected set of operatingconditions of temperature, pressure, and reactant and reaction productflow rates, and the catalysts in the catalyst beds of the stages of saidsecond laboratory scale reactor having selected sets of characteristics;sampling the effluents of each of the reactor stages of said secondmultistage series-connected reactor; measuring the concentration ofproducts and byproducts of the catalytic reactions taking place in saidsecond multistage series-connected reactor in the effluents of each ofsaid reactor stages of said second multistage series connected reactor;repeating steps (f) through (h) at different selected sets of saidoperating conditions, and/or at different selected sets ofcharacteristics of the catalysts in the catalyst beds of said secondlaboratory scale reactor; and using the results of said measurementsobtained in one polynuclear aromatic molecule saturation operation toinfluence the selection of catalyst bed characteristics and operatingparameters in a subsequent polynuclear aromatic molecule saturationoperation for improving the productivity and selectivity of said secondlaboratory scale plug-flow reactor.
 4. The method of claim 3 furtherincluding the steps of: cleaving carbon-carbon bonds in cyclic moleculesin the effluent of the said second laboratory scale reactor by feedingselected partial pressures of effluent from the last stage of saidsecond multistage reactor and hydrogen to the inlet of the first reactorstage of a third multistage series-connected laboratory scale plug flowreactor in which each of the reactor stages includes a catalyst bedcontaining catalyst particles capable of catalyzing the cleaving byhydrogen of carbon-carbon bonds in cyclic molecules, said carbon-carbonbond cleaving process being implemented at a selected set of operatingconditions of temperature, pressure, and reactant and reaction productflow rates, and the catalysts in the catalyst beds of the stages of saidthird laboratory scale reactor having selected sets of characteristics;sampling the effluents of each of the reactor stages of said thirdmultistage series-connected reactor; measuring the concentration ofproducts and byproducts of the catalytic reactions taking place in saidthird multistage series-connected reactor in the effluents of each ofsaid reactor stages of said third multistage series connected reactorrepeating steps (k) through (m) at different selected sets of saidoperating conditions, and/or at different selected sets ofcharacteristics of the catalysts in the catalyst beds of the stages ofsaid third laboratory scale reactor; and using the results of saidmeasurements obtained in one carbon-carbon bond cleaving operation toinfluence the selection of catalyst bed characteristics and operatingparameters in a subsequent carbon-carbon bond cleaving operation forimproving the productivity and selectivity of the third laboratory scaleplug-flow reactor.
 5. The method of claim 4 further including the stepsof: saturating unsaturated molecules in the effluent of said thirdmultistage laboratory scale reactor by feeding selected partialpressures of effluent from the last stage of said third laboratory scalereactor and hydrogen to the inlet of the first reactor stage of a fourthmultistage series-connected laboratory scale plug flow reactor in whicheach of the reactor stages includes a catalyst bed containing catalystparticles capable of catalyzing the saturation by hydrogen ofunsaturated molecules, said unsaturated molecule saturating processbeing implemented at a selected set of operating conditions oftemperature, pressure, and reactant and reaction product flow rates, andthe catalysts in the catalyst beds of the stages of said fourthlaboratory scale reactor having selected sets of characteristics;sampling the effluents of each of the reactor stages of said fourthmultistage series-connected reactor; and measuring the concentration ofproducts and byproducts of the catalytic reactions taking place in saidfourth multistage series-connected reactor in the effluents of each ofsaid reactor stages of said fourth multistage series connected reactorrepeating steps (p) through (r) at different selected sets of saidoperating conditions, and/or at different selected sets ofcharacteristics of the catalysts in the catalyst beds of the stages ofsaid fourth laboratory scale reactor; and using the results of saidmeasurements obtained in one unsaturated molecule saturating operationto influence the selection of catalyst bed characteristics and operatingparameters in a subsequent unsaturated molecule saturating operation forimproving the productivity and selectivity of the fourth laboratoryscale plug-flow reactor.
 6. The method of claim 4 further including thesteps of: hydrocracking and isomerizing hydrocarbon molecules in theeffluent of said third reactor by feeding selected partial pressures ofeffluent from the last stage of said third series-connected reactor andhydrogen to the inlet of the first reactor stage of a fourth multistageseries-connected laboratory scale plug flow reactor in which each of thereactor stages includes a catalyst bed containing catalyst particlescapable of catalyzing the hydrocracking and isomerization by hydrogen ofhydrocarbon molecules, said hydrocracking and isomerization processbeing implemented at a selected set of operating conditions oftemperature, pressure, and reactant and reaction product flow rates, andthe catalysts in the catalyst beds of the stages of said fourthlaboratory scale reactor having selected sets of characteristics;sampling the effluents of each of the reactor stages of said fourthmultistage series-connected reactor; measuring the concentration ofproducts and byproducts of the catalytic reactions taking place in saidfourth multistage series-connected reactor in the effluents of each ofsaid reactor stages of said fourth multistage series connected reactor;repeating steps (u) through (w) at different selected sets of saidoperating conditions, and/or at different selected sets ofcharacteristics of the catalysts in the catalyst beds of the stages ofsaid fourth laboratory scale reactor; and using the results of saidmeasurements obtained in one hydrocracking and isomerization operationto influence the selection of catalyst bed characteristics and operatingparameters in a subsequent hydrocracking and isomerization operation forimproving the productivity and selectivity of the fourth laboratoryscale plug-flow reactor.
 7. The method of claim 5 further including thesteps of: hydrocracking and isomerizing hydrocarbon molecules in theeffluent of said fourth reactor by feeding selected partial pressures ofeffluent from the last stage of said fourth multistage reactor andhydrogen to the inlet of the first reactor stage of a fifth multistageseries-connected laboratory scale plug flow reactor in which each of thereactor stages includes a catalyst bed containing catalyst particlescapable of catalyzing the hydrocracking and isomerization by hydrogen ofhydrocarbon molecules, said hydrocracking and isomerization processbeing implemented at a selected set of operating conditions oftemperature, pressure, and reactant and reaction product flow rates, andthe catalysts in the catalyst beds of the stages of said fifthlaboratory scale reactor having selected sets of characteristics;sampling the effluents of each of the reactor stages of said fifthmultistage series-connected reactor; measuring the concentration ofproducts and byproducts of the catalytic reactions taking place in saidfifth multistage series-connected reactor in the effluents of each ofsaid reactor stages of said fifth multistage series connected reactor;repeating steps (z) through (bb) at different selected sets of saidoperating conditions, and/or at different selected sets ofcharacteristics of the catalysts in the catalyst beds of the stages ofsaid fifth laboratory scale reactor; and using the results of saidmeasurements obtained in one hydrocracking and isomerization operationto influence the selection of catalyst bed characteristics and operatingparameters in a subsequent hydrocracking and isomerization operation forimproving the productivity and selectivity of the fifth laboratory scalereactor.